KR101462049B1 - Process for preparing acrolein or acrylic acid or a mixture thereof from propane - Google Patents

Abstract

The present invention relates to a process for the removal of any water present in product gas A formed in reaction zone A by partially dehydrogenating propane in the reaction zone A under heterogeneous catalysis and at least partially burning the hydrogen molecules formed, , Alternatively using acrolein or acrylic acid or a mixture thereof as the target product from propane which is used to charge reaction zone B, where product gas A is partially oxidized to the target product in the presence of the remaining propane, . The target product is removed from the product gas B formed in the reaction zone B and the propane present in the residual gas remaining is absorbed by the solvent and recycled to the reaction zone A after release from the absorber.

Propane, propylene, acrolein, acrylic acid, reaction zone, recycle

Description

FIELD OF THE INVENTION [0001] The present invention relates to a process for preparing acrolein or acrylic acid or a mixture thereof from propane,

The present invention

A) -propane is fed to the first reaction zone A to form a reaction gas A,

Transferring the reaction gas A through one or more catalyst layers forming hydrogen molecules and propylene by partial heterogeneously catalyzed dehydrogenation of propane in reaction zone A,

Oxygen molecules are supplied to the reaction zone A, at least a part of the hydrogen molecules present in the reaction gas A in the reaction zone A are oxidized with water vapor,

- product gas A containing propylene, propane and water vapor is withdrawn from reaction zone A,

B) Where appropriate, in the first separation zone I, the water vapor present in the product gas A is partially or completely removed by indirect and / or direct cooling of the product gas A to leave the product gas A ^*

C) In reaction zone B, product gas A or product gas A ^* is used together with the supply of oxygen molecules to charge at least one oxidation reactor with reaction gas B comprising propane, propylene and oxygen molecules, Heterogeneously catalyzed partial gas phase oxidation of the propylene to produce acrolein or acrylic acid or a mixture thereof as the target product and also the unconverted propane-containing product gas B,

D) delivering the product gas B out of the reaction zone B and removing the target product present therein in the second separation zone II leaving a propane-

E) If appropriate, a portion of the residual gas having a composition of the residual gas is recycled to the reaction zone A as a propane-

F) In the separation zone III, any water vapor which is not recycled to the reaction zone A, if appropriate, and which is present inside, if appropriate, is partially or completely removed by condensation and / or any hydrogen molecules present therein are pre- (Absorber) by absorption from the residual gas to form a propane-containing absorbent, wherein the propane-

G) In the separation zone IV, the propane is removed from the absorption and recycled to the reaction zone A as the propane-containing feed stream

To a process for the production of acrolein or acrylic acid or mixtures thereof from propane.

Acrylic acid is an important basic chemical substance which provides a use as a monomer for producing a polymer, for example, dispersed in an aqueous medium and used as a binder, in particular. Additional applications of acrylic acid polymers are superabsorbents in the hygiene sector and other applications. Acrolein is an important intermediate for the production of, for example, glutaraldehyde, methionine, 1,3-propanediol, 3-picoline, folic acid and acrylic acid (see, for example, US-A 6,166,263 and US-A 6,187,963).

In a first reaction zone, propane is partially dehydrogenated with propylene under heterogeneous catalysis and the formed propylene is reacted under heterogeneous catalysis in the presence of unconverted propane (i. E., For example, in EP-A 1 617 931 A process for producing acrolein or acrylic acid or mixtures thereof from propane, which partially oxidizes acrolein or acrylic acid or mixtures thereof, without removing propylene from the remaining propane in the pre-dehydrogenation, as recommended in US Pat. (For example, DE-A 33 13 573, EP-A 117 146, US-A 3,161,670, DE-A 10 2004 032 129, EP-A 731 077, DE-A 10 2005 049 699, 10 2005 052 923, WO 01/96271, WO 03/011804, WO 03/076370, WO 01/96270, DE-A 10 2005 009 891, DE-A 10 2005 013 039, DE-A 10 2005 022 798, DE -A 10 2005 009 885, DE-A 10 2005 010 111, DE-A 102 455 85, DE-A 103 160 39, DE 10 2005 056377.5, DE 10 2005 049699.7, DE 10 2006 015235.2, DE 10 2005 061626.7 and prior art cited in these documents).

A common feature of the process described in this document is that after removing the target product acrolein or acrylic acid or acrolein and acrylic acid from the product gas mixture of propylene partial oxidation, at least a portion of the propane present in the residual gas from the remaining propane- And recycled to the heterogeneously catalyzed partial dehydrogenation of propane to propylene as a circulating gas.

Alternatively, the prior art methods can be essentially divided into two groups. One method of either group is to adsorb and remove components other than propane and propylene from the product gas mixture of heterogeneously catalyzed partial dehydrogenation prior to partial oxidation of the propylene to the desired target product in the presence of the remaining propane (DE 10 2005 013039.9). A disadvantage of this procedure is that it generally requires more than two zones to compress the gas. First, it is typically necessary to compress the product gas of inhomogeneously catalyzed partial dehydrogenation of propane prior to removal of the adsorbed propylene and propane present therein, since the absorption elimination can usually only be carried out efficiently under pressure. Also, the compression of the oxygen source used in the heterogeneously catalyzed partial oxidation is generally required, and also the compression of propane, which usually remains in the heterogeneously catalyzed partial oxidation of propylene and recycled to the heterogeneously catalyzed partial dehydrogenation . The other of the two groups is characterized by the use of the product gas mixture itself of heterogeneously catalyzed partial dehydrogenation for the heterogeneously catalyzed partial oxidation of propylene (for example DE-A 103 160 39). The basis of such a procedure is, in particular, the hypothesis that hydrogen molecules formed upon heterogeneously catalyzed partial dehydrogenation of propane do not adversely affect the subsequent propylene heterogeneously catalyzed partial oxidation (requiring the presence of oxygen, (Hydrogen released from hydrocarbons to be dehydrogenated is released directly as water (H ₂ O)). In contrast to pyrogenic heterogeneously catalyzed oxydehydrogenation, which is not detected, heterogeneously catalyzed dehydrogenation (Which will be carried out in reaction zone A) is here an endothermic thermal characteristic as opposed to oxydehydrogenation (pyrogenic hydrogen combustion may be included as a step later in reaction zone A) and free hydrogen molecules are formed at least as intermediates ("Conventional") dehydrogenation, which generally means that different reaction conditions and different catalysts than oxydehydrogenation , I.e. the heterogeneously catalyzed partial dehydrogenation of propane in reaction zone A proceeds with the release of H ₂ in the process according to the invention. However, in-house studies have shown that the presence of increasing amounts of hydrogen molecules in the reaction gas mixture for heterogeneously catalyzed partial oxidation of propylene does not significantly affect the by-product range of partial oxidation (presumably due to the remarkable reduction potential of hydrogen molecules) Lt; RTI ID = 0.0 > catalyst < / RTI > to be used for heterogeneously catalyzed partial gas phase oxidation.

In addition, the hydrogen molecule imparts a potential explosive component with a specific diffusion behavior to the reaction gas B (a particular constituent material is permeable to hydrogen molecules). The use of an excess of oxygen molecules is closely linked to this problem, which is advantageous for increased catalyst life in the partial oxidation of reaction zone B as compared to the target reaction stoichiometry, (When acrylic acid is the target product and the total amount of oxygen molecules required is added to the reaction gas B in advance, it is usually necessary to have a mole ratio of propylene present in the reaction gas B to the oxygen molecule present therein of more than 1.5 Which already takes into account the low degree of propylene complete combustion).

In view of the described prior art, it is an object of the present invention to provide an improved method according to the present disclosure having the disadvantages described in a reduced form even in the worst case.

therefore,

A) -propane is fed to the first reaction zone A to form a reaction gas A,

Transferring the reaction gas A through one or more catalyst layers forming hydrogen molecules and propylene by partial heterogeneously catalyzed dehydrogenation of propane in reaction zone A,

Oxygen molecules are supplied to the reaction zone A, at least a part of the hydrogen molecules present in the reaction gas A in the reaction zone A are oxidized with water vapor,

- product gas A containing propylene, propane and water vapor is withdrawn from reaction zone A,

B) Where appropriate, in the first separation zone I, the water vapor present in the product gas A is partially or completely removed by indirect and / or direct cooling of the product gas A to leave the product gas A ^*

C) In reaction zone B, product gas A or product gas A ^* is used together with the supply of oxygen molecules to charge at least one oxidation reactor with reaction gas B comprising propane, propylene and oxygen molecules, Heterogeneously catalyzed partial gas phase oxidation of the propylene to produce acrolein or acrylic acid or a mixture thereof as the target product and also the unconverted propane-containing product gas B,

D) delivering the product gas B out of the reaction zone B and removing the target product present therein in the second separation zone II leaving a propane-

E) Where appropriate, a portion of the residual gas having a composition of the residual gas is recycled as the propane-containing feed stream to the reaction zone A (residual gas circulation gas)

F) In the separation zone III, any water vapor which is not recycled to the reaction zone A, if appropriate, and which is present inside, if appropriate, is partially or completely removed by condensation and / or any hydrogen molecules present therein are pre- (Absorber) by absorption from the residual gas to form a propane-containing absorbent, wherein the propane-

G) In the separation zone IV, propane is removed from the absorption and recycled to the reaction zone A as a propane-containing feed stream (preferably propane circulating gas containing water vapor)

Wherein at least sufficient hydrogen molecules in the reaction zone A are oxidized with water vapor so that the amount of hydrogen oxidized to water vapor in the reaction zone A is 20 mol% or more of the amount of hydrogen molecules formed in the reaction zone A,

A process for producing acrolein or acrylic acid or mixtures thereof from propane has been found.

Advantageously, according to the invention, at least some of the thermal energy produced by the oxidation (combustion) of hydrogen molecules in the reaction zone A is also used as a heat carrier by indirect heat exchange with the reaction gas A and / or the product gas A The reaction gas A will heat the components of the charge gas mixture (gas feed stream fed to reaction zone A).

Advantageously, according to the invention, the amount of hydrogen oxidized by water vapor in reaction zone A is in each case at least 25 mol%, preferably at least 30 mol%, preferably at least 35 mol%, based on the amount of hydrogen molecules formed in reaction zone A mol% or more, more preferably 40 mol% or more, still more preferably 45 mol% or more, and most preferably 50 mol% or more.

In other words, the process according to the invention can also be carried out in the reaction zone A in an amount of not more than 55 mol%, or not more than 60 mol%, or not more than 65 mol%, or not more than 70 mol%, based on the amount of hydrogen molecules formed in the reaction zone A, Or 75 mol% or less, or 80 mol% or less, or 85 mol% or less, or 90 mol% or less, or 95 mol% or less, or 100 mol% or less.

The fact that the combustion (oxidation) of hydrogen molecules into water provides about twice the amount of heat consumed in the formation of hydrogen molecules during the dehydrogenation of propane is attributed to the fact that hydrogen molecules as components of reaction gas B , The amount of hydrogen molecules burned in the reaction zone A in the process according to the invention is 30 to 70 mol%, preferably 40 to 60 mol%, based on the amount of hydrogen molecules formed in the reaction zone A The case is advantageous in terms of expected heat.

In addition, a limited amount of hydrogen molecules in the reaction gas B (i.e., the reaction gas B still containing hydrogen molecules) is advantageous to some extent in that the hydrogen molecule has the advantage of maximum thermal conductivity in the gas. According to Walter J. Moore, Physikalische Chemie [Physical chemistry], WDEG Verlag, Berlin (1973), page 171, thermal conductivity under standard conditions is, for example, greater than 10 times the thermal conductivity of carbon dioxide and nitrogen molecules or oxygen It is about 8 times the thermal conductivity of the molecule.

This increased thermal conductivity causes the selectivity of CO _x formation in the presence of hydrogen molecules in reaction zone B to be significantly lower in the absence of hydrogen molecules in accordance with Table I of DE-A 33 13 573. Because of the more rapid heat removal from the reaction site in reaction zone B, a lower catalyst surface temperature and hence a lower proportion of propylene complete combustion is caused. This applies in particular when hydrogen molecules with the highest thermal conductivity and propane with the maximum heat absorbency, as in the process according to the invention, are synergistically involved in the heat removal together.

In principle, however, in the process according to the invention, the reaction gas B may no longer contain hydrogen molecules.

In principle, product gas A containing propylene, propane and water vapor and exiting reaction zone A can be used as such in itself to fill one or more oxidation reactors in reaction zone B. However, according to the present invention, at least a portion of the water vapor present in the product gas A in the first separation zone I is condensed by direct and / or indirect cooling of the product gas A, resulting in the removal of water vapor from the product gas A, It is advantageous to leave the product gas A ^* lower than the water vapor content of the product gas A.

One reason for this is that the acrolein and / or acrylic acid target product has a significantly higher affinity than the components of product gas A, which means that product product removal from product gas B, which contains an increased amount of water vapor, This is the reason related to the demand. Secondly, an increased amount of water vapor in the reaction gas B typically contains Mo in oxidized form and reduces the lifetime of the catalyst used in the _secondary oxidation because H ₂ O promotes the sublimation of molybdenum oxide . In addition, the water vapor removed between the reaction zone A and the reaction zone B reduces the total volume of the reaction gas B, thereby reducing the energy requirement to be spent in transferring the reaction gas B through the reaction zone B.

Advantageously, at least 5 mol%, preferably at least 10 mol%, more preferably at least 15 mol%, particularly advantageously at least 20 mol%, more advantageously at least 25 mol% of water vapor present in the product gas A Or at least 30 mol%, even more advantageously at least 35 mol% or at least 40 mol%, preferably at least 45 mol% or at least 50 mol%, particularly advantageously at least 55 mol% or at least 60 mol% Is at least 65 mol% or at least 70 mol%, more preferably at least 75 mol% or at least 80 mol%, even more preferably at least 85 mol% or at least 90 mol%, in many cases at least 95 mol% (But frequently less than or equal to 80 mol%, or less than or equal to 85 mol%, or less than or equal to 90 mol%, or less than or equal to 95 mol%, or less than or equal to 98 mol%) is present in the separation zone I in the process according to the present invention . It is preferable that the limited (but not vanish) water vapor content of the reaction gas B has an advantageous effect on the activity of the catalyst normally used in the reaction zone B (especially in the case of a multimetal oxide).

In the process according to the invention it is preferred that the removal of water between the reaction zone A and the reaction zone B can be achieved by simple condensation.

The water removed between reaction zone A and reaction zone B may not be water formed by hydrogen oxidation only in reaction zone A. Instead, steam may be further used as an inert diluent gas in the reaction zone A and, if appropriate, partially or completely removed in the separation zone I as appropriate.

Because of the relatively high boiling point of the water, especially at elevated pressures, the condensation can usually be achieved by simple direct and / or indirect cooling of the product gas A. In principle, condensation can occur by a single step or by multistage cooling. Suitably according to the present invention, the removal of condensation of water will occur by multi-stage cooling. Advantageously from an application point of view, direct and indirect cooling will be used in combination. Indirect heat exchangers suitable for the purposes of the present invention are, in principle, all known types of indirect heat exchangers. Among them, it is preferable to use a tube heat exchanger and a plate heat exchanger. For example, in the first cooling step, the product gas A discharged from the reaction zone A at a high temperature (the typical temperature of the product gas A is, for example, 400 to 700 占 폚, preferably 500 to 600 占 폚, The working pressure is generally greater than 1 or 1.5 to 5 bar, or 2 to 5 bar, preferably 1.5 or 2 to 3 bar) can be delivered in the first indirect heat exchanger (very generally from reaction zone A The discharged high temperature product gas A may be used in the process according to the invention to heat (in each case, to a desired temperature) at least some of the components or some of the components of the reaction gas A charge gas mixture, while at the same time, (Propane circulating gas), which contains propane and is removed from the absorption in the separation zone IV, in a cocurrent or countercurrent, and, if appropriate, a series of heat exchangers , The product gas A can be brought to the required inlet temperature level in the reaction zone A as one of the propane-containing feed streams delivered to the reaction zone A . This temperature level may vary in a typical manner in the process according to the invention, for example in the range of 400 to 600 占 폚, preferably 450 to 550 占 폚. In the second indirect heat exchanger downstream, the cooled product gas A, as described, is suitably delivered, for example, in co-current or countercurrent to the product gas A which has already been pre-removed to the desired extent, Will be returned to the gas A ^* in either cocurrent or countercurrent flow), the temperature will again increase to the appropriate temperature level and suitable for subsequent partial oxidation of the propylene present therein. When leaving the second indirect heat exchanger, the product gas A generally still has a temperature above 100 ° C. In a third indirect heat exchanger, which is normally installed as an air cooler (e.g., with product gas A in cocurrent or countercurrent to air), the temperature of product gas A will be lowered to a temperature below 100 ° C. Subsequently, the removal of water vapor in a subsequent direct cooler can be carried out. The direct cooler (direct condenser) may have a design known per se (which may correspond to a rectifying column, for example) and may be designed for a purpose that is conventional for the purpose of enlarging the heat exchange surface area ) Internal water. In general, however, such an interior is not necessary. Typically, the column body of the direct condenser is not insulated. Any action that increases the heat loss through the column wall, for example a cooling rib welded in parallel to the perforated sheet or column surface, is preferred. The coolant used for the direct cooling of the product gas A is advantageously cooled to 35 ° C or lower, preferably 30 ° C or lower, in an indirect heat exchanger using cooling water (suitably surface water from an application point of view) Lt; RTI ID = 0.0 > A < / RTI > In the direct cooler, the product gas A and the fine atomized aqueous direct coolant can be delivered in cocurrent or countercurrent. Generally, cocurrent flow is preferable.

Using a separator known per se (e.g., a mechanical drop separator), the aqueous phase and the remaining gaseous phase can be separated from each other. After cooling the removed aqueous phase, which is carried out as described above, the aqueous phase may again be fed to the direct cooling of the product gas A.

The temperature of the remaining gaseous phase (if present, product gas A ^* ) may be increased to a temperature suitable for subsequent partial oxidation of propylene by subsequent indirect heat exchange with product gas A (described above) so that there is no steam. In principle, the water vapor present in the product gas A in the process according to the invention may be removed only by condensation from the part. The resulting product gas A ^* is then subjected to a heterogeneously catalyzed partial oxidation of propylene to a mixture of the product gas A portion which is not treated by itself or by condensation (in a sense, the mixture is likewise the product gas A ^* Lt; / RTI > An additional advantage of direct cooling of the product gas A (for example by use of an aqueous phase which has been pre-removed, preferably by condensation from the product gas A cooled by indirect heat exchange) is generally achieved by direct cooling, Heterogeneously catalyzed partial dehydrogenation, resulting in the deposition of any solid particles present in product gas A, which can then be subjected to heterogeneous catalysis of propylene according to the teaching of DE-A 103 160 39 It can have a cumbersome effect on partial oxidation. This applies in particular when the product gas A or product gas A ^* used to charge the partial oxidation still contains hydrogen molecules. That is to say, according to the present invention, the mechanical separation operation according to DE-A 103 160 39 is very generally carried out in the reaction zone A and the reaction zone B, even if in the form of direct cooling of the reaction gas A, Respectively.

The removal of the condensation of the water vapor from the product gas A in the separation zone I preferably includes the step of cooling the product gas A directly onto the cooled aqueous phase (preferably by indirect heat exchange) Very particularly advantageous.

The aqueous phase removed by condensation, as described, is usually dissolved in a simple manner, for example by heating (if appropriate under reduced pressure) the aqueous phase and by means of a dissolution which can be recovered and recycled to the reaction zone A as an additional propane- ≪ / RTI > propane. Direct and / or indirect cooling of the product gas A for the removal of condensation of water present in the product gas A is generally associated with an internal pressure drop of not more than 1 bar, usually not more than 0.5 bar, preferably not more than 0.3 bar.

Oxidation of hydrogen molecules from reaction zone A to water can be carried out, for example, when a specific partial conversion value is achieved, based on the conversion of propane, which occurs upon single pass of reaction gas A through reaction zone A, Can be carried out after partial dehydrogenation and / or during the process of heterogeneously catalyzed partial dehydrogenation of propane. Advantageously, in accordance with the present invention, at least a portion of the hydrogen oxidized to water in reaction zone A is converted to the desired target conversion (and achieved) of the propane in reaction zone A (based on a single pass of reaction gas A through reaction zone A) Target conversion) is already achieved.

This procedure is advantageous in that it moves the equilibrium position towards the desired propylene formation by, among other things, discharging hydrogen molecules from the equilibrium state of heterogeneously catalyzed dehydrogenation. However, it is further advantageous in that the heat of reaction obtained in the hydrogen oxidation can again replenish the heat of reaction consumed in the pre-performed dehydrogenation and provide a heat of reaction for subsequent further dehydrogenation. Advantageously, according to the invention, the hydrogen molecules (at least a portion) will be burned by the oxygen molecules in reaction zone A until 90 mol%, preferably 75 mol%, of the total amount of hydrogen molecules formed in reaction zone A is formed . Preferably, in accordance with the present invention, the hydrogen molecules (at least a portion) are separated by an oxygen molecule in reaction zone A until 65 mol%, or 55 mol%, or 45 mol% of the total amount of hydrogen molecules formed in reaction zone A is formed It will burn.

The point at which the hydrogen molecule is burned in the reaction zone A can be influenced, among other things, by the location (time) at which the oxygen molecule is fed into the reaction zone A.

When the propane-containing residue gas fraction (the composition of which corresponds to the composition of the residual gas) remaining in the separation zone II is transferred to the reaction zone A on its own (i.e. without passing through the separation zone III / IV) The charge gas mixture in reaction zone A may already contain oxygen molecules since the gas still contains the oxygen molecules used in excess in reaction zone B. [

Also, the composition of the catalyst charge in the reaction zone A can be influenced. Thus, reaction zone A can be locally charged with a catalyst that selectively catalyzes the oxidation of hydrogen molecules to water by exceptionally oxygen molecules. Other catalysts can catalyze dehydrogenation very selectively. In general, catalysts capable of catalyzing dehydrogenation can also catalyze hydrogen combustion, in which case the activation energy of the hydrogen combustion reaction is typically significantly reduced.

Herein, it is understood that the new propane means propane which has not yet participated in dehydrogenation in reaction zone A. Generally, components other than propane are also supplied as minor components of crude propane (preferably meeting the criteria according to DE-A 10246119 and DE-A 10245585).

Such crude propanes are obtainable, for example, by the process described in DE-A 10 2005 022798. In general, in addition to fresh propane, only the additional propane-containing feed stream fed to reaction zone A in the process according to the invention is propane (and, if appropriate, part of the residual gas circulating gas) removed from the absorption in separation zone IV.

In the process according to the invention, it is preferred to exclusively supply fresh propane to reaction zone A as a component of the charge gas mixture for reaction zone A. In principle, the parts of the new propane can also be supplied to the packed gas mixture of the first and / or the second oxidation step of the reaction zone B or any other point in the process for reasons of explosion safety.

In this context, the amount of reactant gas in the catalytic layer catalyzing the reaction step is determined by the amount of standard liters of reactant gas (= 1 (STP)) delivered through one liter of catalyst bed per hour (for example a fixed catalyst bed) (Liters) absorbed under standard conditions (0 DEG C, 1 bar)).

The loading amount can also be based on only one component of the reaction gas. In this case, the loading amount is the amount (l (STP) / l · h) of the component transferred through 1 liter of the catalyst layer per hour (the pure inert material layer is not counted as the fixed catalyst layer). If the catalyst bed consists of a mixture of a catalyst and an inactive molding diluent, the loading may be based on the volume units of catalyst present in reference to this.

In this context, the inert gas generally behaves as a chemically inert material under the corresponding reaction conditions, with greater than 95 mol%, preferably greater than 97 mol%, or 99 mol%, (each inert gas component is considered alone) ≪ / RTI > is understood to mean a reactive gas component that remains chemically unconverted to an extent of greater than < RTI ID = 0.0 >

A useful source for the required oxygen molecules in reaction zone A is, in addition to the residual gas recycle gas derived from separation zone II, a mixture of oxygen molecules themselves or oxygen molecules and gases which chemically inertly react in reaction zone A (E.g., noble gas, such as argon, nitrogen molecules, water vapor, carbon dioxide, etc.). That is, the oxygen molecules are present in an amount of up to 50% by volume, preferably up to 40% by volume or up to 30% by volume, advantageously up to 25% by volume or up to 20% by volume, particularly advantageously up to 15% More preferably not more than 5% by volume, or not more than 2% by volume of other gases (other than oxygen molecules). However, in the process according to the invention, since the oxygen demand in the reaction zone A is relatively low compared to the oxygen demand in the reaction zone B, in the process according to the invention, air is used as the oxygen source for the oxygen demand in reaction zone A It is also desirable, especially because of economic viability. From the viewpoint of the minimum gas volume to be delivered and compressed, it is preferable to feed the reaction zone A as pure oxygen. The propane circulating gas recycled to reaction zone A (from separation zone IV) contains only a small amount of oxygen molecules in general. This is applied in essentially the same way as the supply of oxygen molecules required for reaction zone B. That is, the supply of oxygen molecules as pure oxygen is advantageous to both the reaction zone A and the oxygen demand of the reaction zone B, which is an inert gas that must be re-released during further processing of the circulation process, Because it does not impose on the gas stream.

A typical reaction gas B capable of charging reaction zone B in the process according to the invention has the following contents:

4 to 25% by volume of propylene,

6 to 70% by volume of propane,

0 to 60% by volume of H ₂ O,

O₂ 5 to 60% by volume, and

H₂ 0 to 20% by volume, preferably not more than 10% by volume.

Preferred reaction gases B according to the invention have the following contents:

7 to 25% by volume of propylene,

10 to 40% by volume of propane,

1 to 10% by volume of H ₂ O,

O₂ 10 to 30% by volume, and

H₂ 0 to 10% by volume.

A very particularly preferred reaction gas B according to the invention for said charging has the following contents:

15 to 25% by volume of propylene,

20 to 40% by volume of propane,

H ₂ O 2 to 8% by volume,

O₂ 15 to 30% by volume, and

H₂ 0 to 5% by volume.

The advantage of a suitable propane content in the reaction gas B is evident, for example, from DE-A 102 45 585.

It is preferable that the molar ratio V ₁ of propane present in the reaction gas B to the propane present in the reaction gas B within the above composition range is 1 to 9. It is also advantageous when the molar ratio V ₂ of the propylene present in the reaction gas B to the oxygen molecules present in the reaction gas B within the above composition range is 1 to 2.5. In the context of the present invention, it is also advantageous when the molar ratio V ₃ of propylene present in the reaction gas B to the hydrogen molecules present in the reaction gas B within the above composition range is 0 to 0.80 or 0 to 0.60 or 0.1 to 0.50. Often, V ₃ is also 0.4 to 0.6. It is advantageous when the molar ratio V ₄ of the total molar amount of propane and propylene present in the reaction gas B to the water vapor present in the reaction gas B within the above composition range is 0 or 0.001 to 10.

Particularly advantageously, in the reaction gas B (reaction gas B starting mixture) used to charge reaction zone B, V ₁ is 1 to 7 or 4, or 2 to 6, particularly advantageously 2 to 5, or 3.5 to 4.5 to be. Further, a case where V ₂ is 1.2 to 2.0, or 1.4 or 1.8 is preferable for the reaction gas B starting mixture. Correspondingly, V ₄ in the starting mixture B of the reaction gas B is preferably 0.015 to 5, preferably 0.01 to 3, advantageously 0.01 to 1, particularly advantageously 0.01 to 0.3 or 0.1. A non-explosive reaction gas B starting mixture is preferred in accordance with the present invention.

Reaction gas B An important question to answer as to whether the starting mixture is explosive or not is whether the combustion (ignition, explosion) initiated by a local ignition source (eg, a burning platinum wire) (See experimental description in DIN 51649 and WO 04/007405). When developed, the mixture will be classified as being explosive here. If not developed, the mixture is classified here as non-explosive. When the starting reaction gas mixture of the partial oxidation of the present invention is non-explosive, it is usually also applied to the reaction gas mixture formed during the course of the partial oxidation (see WO 04/007405).

Very generally, the total amount of reaction gas B starting mixture of components other than propylene, hydrogen molecules, water vapor, propane, nitrogen molecules and oxygen molecules is usually 20 vol% or less, or 15 vol% or less, or 10 vol% % Is typical for the process according to the invention. Up to 80% by volume of the other components of this reactant gas B starting mixture may be ethane and / or methane. Alternatively, the components may be in particular carbon oxides (CO ₂ , CO) and rare gases and may also contain minor amounts of secondary component oxygenates such as formaldehyde, benzaldehyde, methacrolein, acetic acid, propionic acid, have. Of course, ethylene, isobutene, n-butane and n-butene are also included in the other possible components of this starting gas B starting mixture. The content of nitrogen molecules in the reaction gas B starting mixture can vary within a wide range. In principle, this nitrogen content can be up to 70% by volume. Generally, the content of nitrogen molecules in the reaction mixture B starting mixture is no more than 65 vol%, alternatively no more than 60 vol%, alternatively no more than 50 vol%, alternatively no more than 40 vol%, alternatively no more than 30 vol%, alternatively no more than 20 vol% 10 vol% or less, or 5 vol% or less. In principle, the N ₂ content of the reaction mixture B starting mixture may be so small as to disappear. Reaction Gas B Minimization of the N ₂ content in the starting mixture minimizes the gas volume to be delivered and compressed in the process according to the invention. Conversely, diluting reaction gas B with nitrogen molecules reduces propionic acid by-product formation in reaction zone B.

In principle, propane-heterogeneously catalyzed partial dehydrogenation processes with all propylene known in the prior art are described, for example, in WO 03/076370, WO 01/96271, EP-A 1 17 146, WO 03/011804, EP A 7 310 77, US 3,162,60, WO 01/96270, DE-A 331 35 73, DE-A 102 45 585, DE-A 103 16 039, DE-A 10 2005 009891, DE-A 10 DE-A 10 2005 029 798, DE-A 10 2005 009 885, DE-A 10 2005 010 111, DE-A 10 2005 049 699, DE-A 10 2004 032 129, DE 10 2005 056377.5, DE 10 2006 017623.5, DE 10 2006 015235.2, DE 10 2005 061626.7 and DE 10 2005 057197.2, as well as in the prior art recognized in these references, for the heterogeneously catalyzed partial dehydrogenation of propane in reaction zone A This is due to the fact that oxidation (combustion) of hydrogen molecules in the reaction zone A by oxygen molecules can in principle be followed by any propane heterogeneously catalyzed partial dehydrogenation process. In principle, the heterogeneously catalyzed dehydrogenation of propane and the combustion of hydrogen molecules in reaction zone A can also be carried out in different reactors, for example in reactors connected in series. In this case, it is advantageous in the process according to the invention that the reaction gas B starting mixture does not necessarily contain oxygen molecules. Very generally, in the reaction zone A, the process according to the invention can be carried out in only one reactor or in more than one reactor connected in series, for example. Based on the single pass of the reaction zone A of the propane fed to the reaction zone A, the reaction zone A is essentially constituted by controlled heat exchange with the fluid (i.e. liquid or gas) heat carrier delivered from outside the reaction zone A . However, it may also be carried out on the same basis, adiabatically, i.e. without a controlled heat exchange with the heat carrier which is essentially transferred outside the reaction zone A. In the latter case, the total thermal properties based on the reaction zone A single pass of the propane fed to reaction zone A are recommended by the literature and can be either endothermic (negative) or self-heating (essentially zero) Or exothermic (positive). It is likewise possible to use all the catalysts recommended in the literature (including the prior art recognized in these documents) in the process according to the invention. In principle, the heterogeneously catalyzed propane dehydrogenation in reaction zone A is carried out adiabatically and / or isothermally (also possible to alternate from adiabatic to isothermal and from isothermal to adiabatic along reaction zone A along reaction zone A) (Or a plurality of fixed bed reactors connected in series, for example) or a moving bed or fluidized bed reactor (the latter is due to backmixing, especially when oxygen molecules have already been fed to the reaction mixture A starting mixture The reaction mixture A is suitable for heating the reaction mixture from the reaction zone A to the reaction temperature by the hydrogen combustion of the reaction gas A).

Typically, propane heterogeneously catalyzed partial dehydrogenation to propylene requires relatively high reaction temperatures. The achievable conversion of propane usually reaches the limit of thermodynamic equilibrium. Typical reaction temperatures are 300 to 800 占 폚, or 400 to 700 占 폚, or 450 to 650 占 폚. One hydrogen molecule per propane molecule dehydrogenated with propylene is produced. According to the invention, the working pressure in the entire reaction zone A is suitably from 1 to 5 bar, preferably from 1.5 to 4 bar, advantageously from 2 to 3 bar. In principle, the working pressure in the reaction zone A may also be a value higher or lower than this value. Removal of the hot and H ₂ reaction products shifts the dehydrogenation equilibrium position in reaction zone A from the reaction zone A to the propylene to be formed.

Since the heterogeneously catalyzed dehydrogenation reaction proceeds with increasing volume, the conversion can be increased by lowering the partial pressure of the dehydrogenation product. This can be done in a simple manner, for example by dehydrogenation under reduced pressure (however, the performance at elevated pressure is generally favorable for the catalyst life) and / or an essentially inert diluent gas constituting an inert gas for the dehydrogenation reaction, Can be achieved by adding water vapor. As a further advantage, dilution with water vapor generally results in a reduction in the coking of the catalyst used, since the water vapor reacts with the carbon formed by the coal gasification principle. The heat capacity of the water vapor can also supplement a portion of the endotherm of the dehydrogenation.

A limited amount of water vapor has usually been found to have a promoting effect on the activity of the partial oxidation catalyst in the downstream reaction zone B, but amounts exceeding this are disadvantageous in reaction zone B for the reasons mentioned above. According to the present invention, at least a portion of the hydrogen must be further oxidized to water vapor in the reaction zone A. Therefore, when the amount of the water vapor supplied to the zone A is 20 vol% or less, preferably 15 vol% or less, more preferably 10 vol% or less, based on the reaction gas A supplied to the charging of the catalyst in the reaction zone A It is advantageous according to the invention. In general, however, the amount of water vapor fed to the reaction zone A will typically be at least 1 vol.%, In many cases at least 2 vol.%, Or at least 3 vol.%, Often at least 5 vol.% On the same basis.

Additional diluents suitable for heterogeneously catalyzed propane dehydrogenation are, for example, nitrogen, noble gases such as He, Ne and Ar, as well as compounds such as CO, CO ₂ , methane and ethane. All the diluents mentioned can be used alone or in the form of different mixtures. If the diluent gas is formed as a by-product in the recycle gas process according to the invention or is supplied as a fresh gas or a new gas component, the process according to the invention requires a corresponding manner of release, Is a less preferred reason according to the present invention. These possible emissions are in different separation zones, especially separation zones III and IV.

In principle, dehydrogenation catalysts useful for heterogeneously catalyzed propane dehydrogenation are all known in the art. They are generally divided into two groups, specifically catalyst groups (e.g. chromium oxide and / or aluminum oxide) and oxide supports (for example zirconium dioxide, magnesium oxide, aluminum oxide, silicon dioxide, titanium dioxide , Lanthanum oxide and / or cerium oxide) deposited on one or more surfaces of the substrate (e.g., platinum). The dehydrogenation catalysts which can be used thus are described in WO 01/96270, EP-A 731077, DE-A 10211275, DE-A 10131297, WO 99/46039, US-A 4 788 371, EP-A 705 136, WO 99/29420 , US-A 4 220 091, US-A 5 430 220, US-A 5 877 369, EP-A 117 146, DE-A 199 37 196, DE-A 199 37 105, -A 6 566 573, WO 94/29021 and DE-A 199 37 107, all of which are incorporated herein by reference. In particular, the catalysts according to Examples 1, 2, 3 and 4 of DE-A 199 37 107 can be used. Finally, the catalyst of German Application No. 10 2005 044916 should also be recommended as a catalyst to be used for dehydrogenation in reaction zone A. In the process according to the invention, these catalysts can be the sole catalysts used in the reaction zone A since they generally catalyze the combustion of hydrogen molecules by oxygen molecules.

In principle, the catalytic active layer in the reaction zone A can be made exclusively of a catalytically active molded body. However, it is of course possible that the catalytically active layer in the reaction zone A consists of the catalytically active molded body diluted with the inert molded body. Such an inert shaped body may be, for example, a calcined clay (aluminum silicate) or a soap (e.g. C 220 of CeramTec), or another (preferably essentially void free) high temperature ceramic material, , Silicon dioxide, titanium dioxide, magnesium oxide, lanthanum oxide, cerium oxide, zinc aluminum mixed oxide, thorium dioxide, zirconium dioxide, silicon carbide or other silicates such as magnesium silicate and mixtures of these materials. This material is also useful for the inert top layer of the fixed catalyst bed of the process according to the invention and, if appropriate, the last layer.

The dehydrogenation catalyst comprises 10 to 99.9% by weight of zirconium dioxide, 0 to 60% by weight of aluminum oxide, silicon dioxide and / or titanium dioxide, and one or more elements of the first or second group of elements in the periodic table of the elements, 0.1 to 10% by weight of tantalum, one element of the eighth transition group, lanthanum and / or tin (assuming that the sum of the weight% is less than 100% by weight).

Also, the dehydrogenation catalysts used in the examples herein are particularly suitable.

In a preferred embodiment, the dehydrogenation catalyst comprises one or more elements of the parent family VIII, one or more elements of the families I and II, one or more elements of the families III and / or IV, (Including lanthanide and actinide elements). As an element of total bovine VIII, the active composition of the dehydrogenation catalyst preferably comprises platinum and / or palladium, more preferably platinum. As elements of families I and II, the active composition of the dehydrogenation catalyst preferably comprises potassium and / or cesium. Lanthanide and actinide elements, wherein the active composition of the dehydrogenation catalyst preferably comprises lanthanum and / or cerium. As the elements of families III and / or IV, the active composition of the dehydrogenation catalyst preferably comprises at least one element selected from the group consisting of boron, gallium, silicon, germanium, tin and lead, more preferably tin . Most preferably, the active composition of the dehydrogenation catalyst comprises at least one representative element in each case from the element family.

Typically, the dehydrogenation catalyst is a catalyst extrudate (typically 0.1 or 1 to 10 mm in diameter, preferably 1.5 to 5 mm in diameter, typically 1 to 20 mm in length, preferably 3 to 10 mm in diameter) Tablets (preferably of the same dimensions as for extrudates) and / or catalytic rings (in each case having an outer diameter and a length of typically 2 to 30 mm or 10 mm, a wall thickness of preferably 1 to 10 mm, or 5 mm , Or 3 mm). To effect heterogeneous catalytic dehydrogenation in the fluidized bed (or mobile bed), correspondingly finer catalysts will be used. For the reaction zone A, it is preferred to use a fixed catalyst bed according to the invention.

In principle, there are no limitations regarding the shape of the catalyst to be used for the fixed catalyst bed (in particular for supported catalysts) or the inert molded body shape to be used in the process according to the invention. Particularly often used shapes are solid cylinders, hollow cylinders, spheres, cones, pyramids and cubes, and also extrudates, wagonwheels, stars and monoliths.

The longest dimension of the catalyst formed body and the inert molded body (the longest straight line connecting the two points on the molded body surface) may be 0.5 mm to 100 mm, often 1.5 mm to 80 mm, in many cases 3 mm to 50 mm or 20 mm.

In general, dehydrogenation catalysts (in particular those used herein, and those recommended in DE-A 19 937 107, in particular the exemplary catalysts in this German application), are used for the dehydrogenation of propane and hydrogen molecules and propane / So as to catalyze both of the combustion of the fuel. Hydrogen combustion proceeds very rapidly compared to both dehydrogenation of propane and, for example, combustion of propane in the case of a competition situation on the catalyst (i.e. the catalyst has the lowest activation energy for the combustion of hydrogen molecules under certain conditions ). Combustion of propane / propylene generally proceeds more rapidly than dehydrogenation on the catalyst.

Said coupling enables the use of only one catalyst type in reaction zone A in the process according to the invention. Advantageously, in the reaction zone A, when the reaction gas A comprises oxygen molecules and is in contact with the catalyst described above, the reaction gas A contains an amount of hydrogen molecules that is at least twice the amount of oxygen molecules present therein This is because there is a risk of combustion of propane and / or propylene and the yield of the target product is reduced based on the converted propane feedstock in the process according to the present invention.

Of course, the catalyst used to burn hydrogen molecules in water in reaction zone A may also be to specifically catalyze this reaction. Such catalysts are disclosed, for example, in documents US-A 4788371, US-A 4886928, US-A 5430209, US-A 5530171, US-A 5527979 and US-A 5563314.

To perform heterogeneously catalyzed propane dehydrogenation, useful reactor types and process variants are in principle all known in the art. The description of such process variants is present in all prior art documents cited, for example, regarding dehydrogenation catalysts and also in the prior art cited at the beginning of this document.

A relatively comprehensive description of a suitable dehydrogenation process in accordance with the present invention is also provided in Catalytica ^® Studies Division, Oxidative Dehydrogenation and Alternative Dehydrogenation Process, Study Number 4192 OD, 1993, 430 Ferguson Drive, Mountain View, California, 94043-5272 USA exist.

Due to the relatively high reaction temperature required for heterogeneously catalyzed dehydrogenation of propane, small amounts of high boiling point high molecular weight organic compounds (including carbon) are generally formed in the reaction zone A in the process according to the invention and deposited on the catalyst surface It becomes practical. In order to minimize this adverse co-occurrence, the propane-containing reaction gas A which will pass over the catalyst surface at elevated temperatures for heterogeneously catalyzed dehydrogenation can be suitably diluted with water vapor. This is carried out in a manner particularly advantageous in the process according to the invention, for example, by stripping with water vapor or steam-containing gas (for example an inert gas) to effect removal of the propane from the absorbent in the separation zone IV and to remove the propane- (Propane-containing water vapor) into the reaction zone A as propane circulating gas. The steam will partially or completely remove the carbon deposited by the principle of coal gasification under otherwise existing conditions to prolong the catalyst life.

Another method of removing the deposited carbon compounds is by flowing an oxygen-containing gas through the dehydrogenation catalyst (and optionally any catalyst separately used for hydrogen combustion) at elevated temperatures (suitably in the absence of hydrocarbons) And effectively burning and removing carbon. However, the inhibition of certain carbon deposit formation and thus the catalyst lifetime extension is also possible by adding hydrogen molecules to the propane to be dehydrogenated under heterogeneous catalysis before being passed over the dehydrogenation catalyst at elevated temperatures. This is because in the process according to the invention any hydrogen molecules still present in the residual gas, suitably present in the separation zone II, for example, are removed from the separation zone II by the separation membrane III (or beyond the separation zone III) It is possible in a simple manner to recycle to Zone A. Of course, it is also possible to use hydrogen molecules from another source at this time. For example, the hydrogen molecule may also still suitably be present in the residual gas recycle gas recycled to reaction zone A.

Of course, there is also the possibility of adding a mixture of water vapor and hydrogen molecules to the propane to be dehydrogenated under heterogeneous catalysis.

The addition of hydrogen molecules to heterogeneously catalyzed dehydrogenation of propane also reduces the formation of unwanted allenes (propadiene), propyne and acetylene as byproducts.

This allows, in the simplest case, only new propane and propane circulating gases to be fed into reaction zone A to form reaction gas A to be delivered through one or more catalyst beds (here, the charge gas mixture of reaction zone A or reaction gas A starting mixture A Lt; / RTI > The propane circulating gas is used in accordance with the present invention to remove oxygen molecules in an amount required to cause the required hydrogen combustion in the reaction zone A (e.g., propane is stripped off from the absorbent in the separation zone IV and the stripping gas- For example, in the form of air). However, the reaction gas A starting mixture may also simply consist of fresh propane, propane circulating gas and residual gas circulating gas, in which case the latter may likewise contain oxygen molecules. In an advantageous manner, the propane recycle gas in this case will contain sufficient water vapor to develop favorable properties for the overall process together with the water vapor formed during the course of hydrogen combustion in reaction zone A. In the case described above, it is not necessary to supply additional gas stream to reaction zone A. The desired conversion in reaction zone A proceeds upon a single pass of reaction gas A.

Of course, the reaction gas A to be conveyed through one or more stationary catalyst layers may contain additional water vapor and / or additional hydrogen molecules in addition to the new propane, propane circulation gas and, preferably, residual gas circulation gas to develop the advantageous function described herein . The molar ratio of hydrogen molecules to propane in the packed gas mixture in reaction zone A is generally below 5, usually below 3, often below 1 or below 0.1.

The molar ratio of steam to propane in the charge gas mixture in reaction zone A may be, for example, from 0 to 30. Suitably 0.05 to 2, preferably 0.1 to 0.5.

It is also possible to add additional oxygen molecules (mixture with pure form and / or inert gas) and / or additional inert gas as needed for the fill gas mixture for reaction zone A. The desired conversion in reaction zone A proceeds again in the single pass of reaction gas A (the charge gas in reaction zone A) without the additional gas stream being fed along the reaction path. As used herein, the reaction path in reaction zone A is a function of the conversion of dehydrogenation (conversion in heterogeneously catalyzed dehydrogenation) of propane fed to reaction gas A prior to single pass through one or more catalyst beds in reaction zone A Is understood to mean the flow path of propane through reaction zone A.

Suitable reactor configurations for the heterogeneously catalyzed propane dehydrogenation, in which the packed gas mixture passes through the reaction zone A without intermediate gas feed, are, for example, fixed bed tubular reactors or vortex reactors. In this reactor, the dehydrogenation catalyst is disposed as a fixed bed in one reaction tube or reaction tube bundle. In accordance with the present invention, if the total hydrogenation of the hydrogen required in the reaction zone A is such that the total reaction proceeding in the reaction zone A is endothermic, the reaction tube will suitably be heated externally according to the invention Will be understood). This can be done, for example, by burning hydrocarbons such as gas, e.g., methane, in the space around the reaction tube. This direct form of catalyst tube heating is only applied to the first 20-30% of the fixed bed and the remaining bed length is preferably heated to the required reaction temperature through the radiant heat released during the combustion process. In this way, approximately isothermal reaction is achievable. Suitable reaction tube diameters are about 10 to 15 cm. Typical dehydrogenation vortex reactors include from 300 to more than 1000 reaction tubes. The temperature inside the reaction tube varies within the range of 300 to 800 占 폚, preferably 400 to 700 占 폚. Advantageously, the reaction mixture A is fed to a tubular reactor preheated to the reaction temperature. The product gas (mixture) A is able to escape from the reaction tube having a lower temperature of 50 to 100 ° C. However, this outlet temperature may be at a higher level or at the same level. In this procedure, the use of an oxidative dehydrogenation catalyst based on chromium oxide and / or aluminum oxide is suitable. The dehydrogenation catalyst is usually used without dilution. On an industrial scale, a plurality of said vortex reactors may be operated in parallel, and product gas A may be used to charge reaction zone B as a mixture. Suitably, it is also possible that only two of these reactors are involved in the dehydrogenation operation and the catalyst charge is regenerated in the third reactor.

A single pass of the packed gas through the reaction zone A can also be carried out in a moving bed or fluidized bed reactor, for example as described in DE-A 102 45 585 and in the literature cited for this topic in this document.

In principle, the reaction zone A of the process according to the invention can also for example consist of two compartments. The structure of this reaction zone A is particularly valid when the charge gas for reaction zone A does not contain any oxygen molecules.

In this case, inhomogeneous catalytic dehydrogenation in practice can be carried out in the first compartment and, after immediate supply of a mixture of oxygen molecules and / or oxygen molecules and an inert gas, the heterogeneously catalyzed hydrogen combustion required according to the invention Can be performed in the second compartment.

Very generally, in the process according to the invention, reaction zone A is suitably not less than 5 mol% to not more than 60 mol% of the total amount of propane fed to reaction zone A, Is not less than 10 mol% to not more than 50 mol%, more preferably not less than 15 mol% to not more than 40 mol%, most preferably not less than 20 mol% to not more than 35 mol% Will be. This limited conversion rate in reaction zone A is usually sufficient in the process according to the invention because the remaining amount of unconverted propane functions as substantially diluent gas downstream of reaction zone B and substantially during the course of the process of the invention Because it can be recycled without loss to Zone A. An advantage of the procedure with a relatively low propane conversion rate is that the amount of heat required for the endothermic dehydrogenation during the single pass of reaction gas A through reaction zone A is relatively small and a relatively low reaction temperature is sufficient to effect conversion.

Advantageously, according to the invention, it may be appropriate to carry out the propane dehydrogenation in the reaction zone A (sub-adiabatically) (for example with a relatively low propane conversion), as already explained. This means that the charge gas mixture for reaction zone A will generally be heated first to a temperature of 500 to 700 占 폚 (or 550 to 650 占 폚, for example, by direct ignition of the peripheral wall). Usually, the one-time thermal insulting through the catalyst layer is carried out in such a manner that during the course of the reaction gas being heated, cooled, or behaving in a thermal neutrality manner in accordance with the quantitative ratio of endothermic dehydrogenation to exothermic hydrogen combustion, Will be sufficient to achieve both conversion and hydrogen burning required in accordance with the present invention. According to the present invention, an adiabatic operation mode in which the reaction gas is cooled to about 30 to 200 DEG C during single pass is preferable. If necessary, in the second compartment of the reaction zone A, the hydrogen formed during the dehydrogenation can be post-burned into the oxygen molecules fed intermediate in the heterogeneous catalysis. This combustion can be carried out equally adiabatically. Significantly, especially in adiabatic operation, a single shaft furnace which axially and / or radially flows the reaction gas A is sufficient as a fixed bed reactor.

In the simplest case, this is a vessel with a single closed reaction volume, for example an inner diameter of 0.1 to 10 m, possibly even 0.5 to 5 m, and a fixed catalyst bed disposed on a support device (e.g. a grid). The reaction volume charged with the catalyst and substantially insulated in the adiabatic operation flows axially by the hot reaction gas A containing propane. The catalyst shape may be spherical or annular or extruded. In this case, all catalyst shapes with a particularly low pressure drop are preferred, since the reaction volume can be realized by a very inexpensive device. These are in particular a catalytic form or a structural form, for example a monolith or a honeycomb, which leads to a large cavity volume. In order to realize the radial flow of the reaction gas A containing propane, the reactor may be composed of, for example, two concentric cylinder lattices arranged in the shell, and the catalyst layer may be arranged in the annular gap. In the case of insulation, the metal shell will be insulated if appropriate.

The catalyst packings of the present invention suitable for heterogeneously catalyzed propane dehydrogenation are in particular the catalysts disclosed in DE-A 199 37 107, in particular all those given as examples and mixtures thereof with geometric shaped bodies which are inert towards heterogeneous catalytic dehydrogenation to be.

After prolonged working hours, the catalyst may be heated, for example, in a simple manner by air (preferably) diluted with nitrogen and / or steam onto the catalyst bed in a first regeneration step at a temperature of 300 to 600 占 폚, often 400 to 550 占 폚 And can be regenerated by passing it at an initial temperature at an early stage. The amount of catalyst (layer) loading by the regeneration gas (for example air) may be, for example, 50 to 10 000 h ^-1 and the oxygen content of the regeneration gas may be 0.1 or 0.5 to 20% by volume.

Thereafter, in the additional regeneration step, air can be used as the regeneration gas under otherwise identical regeneration conditions. Suitably from an application point of view, it is advisable to flush the catalyst with an inert gas (e.g. N ₂ ) before regeneration.

It is then reasonable to regenerate hydrogen molecules (hydrogen content should be at least 1% by volume) diluted with pure hydrogen molecules or an inert gas (preferably water vapor and / or nitrogen) generally under otherwise identical conditions.

Heterogeneously catalyzed propane dehydrogenation in the reaction zone A of the process according to the invention is a variant with a high propane conversion (greater than 30 mol%) (with respect to both the reaction gas and all the propane present therein) as the same catalyst (30 mol% or less) in all cases. The loading amount of the reaction gas A may be, for example, 100 to 40000 or 10000 h ^-1 , often 300 to 7000 h ^-1 , that is, in many cases, about 500 to 4000 h ^-1 . The loading amount can also be applied to any catalyst layer used specifically for hydrogen combustion in reaction zone A.

In a particularly preferred manner heterogeneously catalyzed propane dehydrogenation in the reaction zone A (especially in the case of 15 to 35 mol% propane conversion on a single pass basis) can be carried out in a tray reactor, Is the reason for including one or more of the above reactors.

The reactor comprises at least one catalyst bed that catalyzes dehydrogenation in series. The number of catalyst layers may be from 1 to 20, suitably from 2 to 8, and from 3 to 6. Increased propane conversion can be more easily achieved with an increased number of trays. The catalyst layers are preferably arranged radially or axially continuously. Suitably from an application point of view, a fixed catalyst bed type is used for this tray reactor.

In the simplest case, the fixed catalyst bed in the reactor as a shaft is disposed axially or in the annular gap of the concentric cylindrical lattice. It is also possible, however, to arrange the annular gap in a compartmentalized layer, to pass the gas radially through one compartment and into the next compartment above or below it.

Suitably, the reaction gas A is passed through a heat exchanger surface (for example, a lip) heated by a hot gas, for example, from one catalyst layer to a next catalyst layer, or passes through a heated pipe by a hot combustion gas, (If necessary, intermediate cooling in a corresponding manner).

When the tray reactor is operated adiabatically differently, especially when the catalyst described in DE-A 199 37 107, especially the catalyst of the exemplary embodiment, is used, the reaction gas mixture is heated to a temperature of 450 to 550 캜 (preferably 450 to 500 캜) It is sufficient for the propane conversion to be delivered to the preheated dehydrogenation reactor to temperature and to maintain this temperature range within the tray reactor, especially up to 30 mole%. This means that the overall propane dehydrogenation can be realized at very low temperatures, which has been found particularly desirable for the lifetime of the fixed catalyst bed between two regeneration. For higher propane conversion rates, the reaction gas mixture is suitably delivered to the preheated dehydrogenation reactor at a higher temperature (which may be below 700 캜) and is maintained within this temperature range within the tray reactor.

It is even more desirable to perform the above described intermediate heating at the direct method (which allows the self-heating method). To this end, a limited amount of oxygen molecules are introduced into the first catalyst bed (for example as a component of the residual gas circulation gas and / or propane circulation gas), in which case the reaction mixture A must advantageously contain hydrogen molecules ) And / or between the downstream catalyst layers. Thus, the combustion of hydrogen molecules which are required in accordance with the invention and which are present in the reaction gas A and / or which are formed during the process of heterogeneously catalyzed propane dehydrogenation and / or are added in a particularly selective and controlled manner to the reaction gas A (Which is generally catalyzed by the dehydrogenation catalyst itself), which may also be appropriate from an application point of view, in particular also (optionally) inserting the catalyst layer into a tray reactor filled with a catalyst that catalyzes the combustion of hydrogen Examples of useful catalysts are described in US-A 4 788 371, US-A 4 886 928, US-A 5 430 209, US-A 5 530 171, US-A 5 527 979 and US- For example the catalyst bed can be mounted in a tray reactor alternately in a layer comprising a dehydrogenation catalyst, these catalysts are also suitable for the hydrogen combustion in the second compartment of the reaction zone A) . Depending on the amount of hydrogen molecules burned, the emitted heat of reaction may be directed to the entire exothermic operating mode for heterogeneously catalyzed propane dehydrogenation, or the entire self-heating operating mode (where the total thermal property is substantially zero) Enables the whole. As the selected residence time of the reaction gas in the catalyst bed increases, propane dehydrogenation is possible at a reduced or substantially constant temperature, which enables a long catalyst life, especially between two regimes, and is preferred according to the invention.

In particular, when the hydrogen molecule is burned in both the heterogeneously catalyzed dehydrogenation of propane (step 1) and the heterogeneously catalyzed dehydrogenation of propane (step 2), it is preferable to firstly introduce heterogeneously catalyzed dehydrogenation The reaction gas present in the reaction zone A that is present after completion is recycled as the dehydrogenation circulating gas to the reaction zone A in an amount supplied to the final hydrogen combustion (partly as one of the propane-containing feed streams) Gas and dehydrogenation are cooled by indirect heat exchange using the components of the reaction mixture A other than the circulating gas A and the product gas A after the final combustion of the hydrogen molecules in the reaction zone A is likewise subjected to indirect heat exchange It is appropriate to carry out additional treatment after cooling. This applies in particular when the reaction zone A is constituted as a whole exothermic.

Generally, the oxygen feed as described above is to be fed in such a manner that the oxygen content of the reaction gas A is from 0.5 to 50 or 40 or 30% by volume, preferably from 10 to 25% by volume, based on the amount of hydrogen molecules present therein Should be carried out according to the invention. Useful oxygen sources include pure oxygen molecules (preferred in accordance with the invention) or oxygen molecules as diluted with an inert gas such as CO, CO ₂ , N _2, and / or rare gas, . Preferably, according to the present invention, the oxygen molecules are present in an amount of up to 30% by volume, preferably up to 25% by volume, advantageously up to 20% by volume, more advantageously up to 15% by volume, preferably up to 10% And preferably not more than 5 vol% of other gases (other than oxygen molecules). Particularly advantageously, the oxygen supply is carried out in pure form.

As the combustion of 1 mol hydrogen molecules into H ₂ O provides about twice (approximately 240 kJ / mol) of the dehydrogenation energy of propylene and 1 mol propane to H ₂ (approximately 120 kJ / mol) The self-heating configuration of the reaction zone A in the adiabatic tray reactor is effectively combined with the considered advantages of the process of the present invention, where it is only possible to combine the amount of hydrogen produced in dehydrogenation in the reaction zone A with the amount of hydrogen of about 50 mol% ≪ / RTI >

However, an advantage of the process according to the invention is that it is not only obtained when about 50 mol% of the amount of hydrogen formed in reaction zone A is burned in reaction zone A. Alternatively, this advantage may be achieved by adding 25 to 95 mol% or 100 mol%, or 30 to 90 mol%, or 35 to 85 mol%, or 40 to 80 mol%, or 45 to 75 mol% of the amount of hydrogen molecules formed in reaction zone A , or 50 to 70 mol%, alternatively 35 to 65 mol%, or 40 to 60 mol%, or 45 to 55 mol% of the total amount of hydrogen is combusted in the reaction zone A to provide water To the above operating mode of the adiabatic tray reactor).

Generally, the oxygen feed as described above should be performed such that the oxygen content of the reaction gas A is 0.01 or 0.5 to 3% by volume based on the amount of propane and propylene present therein.

The isothermicity of the heterogeneously catalyzed propane dehydrogenation may be further improved by introducing a closed, e.g., tubular, interior material that is preferably, although not necessarily, vacuum treated before filling into the space between the catalyst beds in the tray reactor. Such an inner material may be located in a specific catalyst layer. These internal materials include suitable solids or liquids that evaporate or melt above a certain temperature, thereby consuming heat and re-condensing when the temperature falls below this value, thereby releasing heat.

Another means of heating the packed gas mixture to the requisite reaction temperature for heterogeneously catalyzed propane dehydrogenation in the reaction zone A is to separate the propane and the propane present therein by the oxygen molecules present in the packed gas mixture upon introduction into the reaction zone & / or burn a portion of the H ₂ and to (for the a particular combustion catalyst suitable for containing, for example, that the top and / or a simple by-pass therethrough), the preferred reaction temperature in the dehydrogenation by the emitted heat of combustion thus (This method is advantageous especially for fluidized bed reactors (as already mentioned)).

Preferably, according to the invention, the reaction zone A will be constructed adiabatically (with external insulation).

According to the above, the reaction zone A of the process according to the invention can be configured as described in documents DE-A 10 2004 032 129 and DE-A 10 2005 013 39, Propane, steam-containing propane circulating gas, and (if appropriate, oxygen molecule-containing) residual gas circulating gas. The reaction zone A is implemented as a tray reactor (preferably an adiabatic) in which a catalyst layer (preferably a fixed bed) is disposed in a radial or axial continuum. Advantageously, the number of catalyst bed trays in this tray reactor is three. It is preferred that the heterogeneously catalyzed partial propane dehydrogenation is performed thermally. To this end, a limited amount of the oxygen molecule or mixture comprising the inert gas is introduced into the reaction zone A beyond the first (fixed) catalyst bed passing through the catalyst bed (fixed) Is added to the charge gas mixture (as described, for example, in DE 10 2006 017623.5). Thus, a limited combustion of hydrogen formed during the heterogeneously catalyzed propane dehydrogenation process (and, if appropriate, at least to a minimum extent, the dehydrogenation of propane and / or propylene) catalyzed by the dehydrogenation catalyst itself can occur, where The pyrogenic thermal property essentially maintains the dehydrogenation temperature.

Suitably, the partially heterogeneous catalyzed dehydrogenation of the propane is carried out essentially on three catalyst trays so that the conversion of propane delivered into the reactor is about 20 mol% based on the passage of a single reactor (method according to the invention Or 30 mole%, or 40 mole%, or 50 mole% or more). The selectivity of the achieved propylene formation is regularly 90 mol%. The maximum contribution of a single tray to the transition moves from front to back in the flow direction with increasing working time. Generally, the catalyst charge is regenerated before the third tray provides the maximum contribution to the transition in the flow direction. Advantageously, regeneration is performed when carbonization of all trays is achieved to the same degree.

Finally, in the (fixed) catalyst bed disposed downstream of the dehydrogenation, after preliminary intermediate oxygen feed, essentially combustion of hydrogen molecules can also be performed. This can in principle proceed to such an extent that the product gas A leaving the reaction zone A is essentially no longer containing any hydrogen molecules.

The heterogeneously catalyzed partial dehydrogenation of propane is carried out in such a way that the loading (total on all layers) to the total amount of catalyst having the total amount of propane and propylene is 500 l (STP) / l · h and 20000 l (STP) h (typical values are 1500 l (STP) / l · h to 2500 l (STP) / l · h). The maximum reaction temperature in the individual fixed catalyst bed is advantageously maintained at 500 ° C to 600 ° C (or 650 ° C).

The disadvantage of the process using a reactant gas A filler mixture already containing oxygen molecules is that virtually all catalysts which catalyze the dehydrogenation of propane are also able to combine the combustion of propane and propylene with oxygen molecules Complete oxidation).

In addition to the action of adding hydrogen molecule-containing gas to the reaction gas A charge gas mixture as described above, it is also possible to carry out the dehydrogenation zone according to DE-A 102 11 275, ), Divided into two parts of the same composition, and only one part of the two parts of the product gas A or the product gas A ^{* (} after the preceding part or the complete combustion of the hydrogen molecules present therein if appropriate) And reacting the other portion with the dehydrogenation as a component of the reaction gas A (generally the reaction gas A packed gas mixture). The hydrogen molecules present in the dehydrogenated recycle gas from the dehydrogenation itself are then intended to protect the propane present in the packed gas mixture for reaction zone A and, if appropriate, propylene from oxygen molecules present therein as well. This protection is based on the fact that the combustion of hydrogen molecules into water, which is heterogeneously catalyzed by the same catalyst as described above, is mechanically preferable to the complete combustion of propane and / or propylene.

According to the teaching of DE-A 102 11 275, the dehydrogenation recycle gas circulation will be realized by suitably the jet-jump principle (also referred to as loop mode: the propane circulation gas can act as a drive jet). The document also teaches the possibility of addition of hydrogen molecules to the charge gas mixture for reaction zone A as further oxidation protection (this may be done, for example, by separation membranes derived from the separation zone II and still from the residual gas comprising hydrogen molecules Which may be a removed hydrogen molecule).

In the process according to the invention, the typical Reaction Gas A fill gas mixture (Reaction Gas A formed in Reaction Zone A) has the following contents:

55 to 80% by volume of propane,

0.1 to 20% by volume of propylene,

H₂ 0 to 10% by volume,

O₂ 0 to 5% by volume,

N₂ 0 to 20% by volume, and

H ₂ O 5 to 15% by volume.

The reaction gas A fill gas mixture (gas mixture introduced into the first fixed catalyst bed in the reaction zone A in the flow direction of the reaction gas A) preferred for the process according to the present invention has the following contents:

55 to 80% by volume of propane,

0.1 to 20% by volume of propylene,

H₂ 2 to 10% by volume,

O₂ 1 to 5% by volume,

N₂ 0 to 20% by volume, and

H ₂ O 5 to 15% by volume.

The proportions of components of the reaction gas A fill gas mixture other than the listed components are generally not more than 10% by volume, preferably not more than 6% by volume.

Typical product gas A has the following contents:

30 to 50% by volume of propane,

15 to 30% by volume of propylene,

H₂ 0 to 10% by volume, preferably 0 to 6% by volume,

10 to 25% by volume of H ₂ O,

O₂ 0 to 1% by volume, and

N₂ 0 to 35% by volume.

The composition ratio of the product gas A other than the listed components is generally not more than 10% by volume. The temperature of the product gas A discharged from the reaction zone A is typically 300 to 800 캜, preferably 400 to 700 캜, more preferably 450 to 650 캜.

The pressure of the product gas A flowing out of the reaction zone A is generally from 1 to 5 bar, preferably from 1.5 to 4 bar, advantageously from 2 to 3 bar.

According to the invention, the product gas A is then introduced into the product gas without removal of further secondary components (but preferably after mechanical removal of the solid particles present therein according to the teaching of DE-A 10 316 039) Is used to charge one or more oxidation reactors in reaction zone B together with reaction gas B after some or all of the condensation removal (i.e., as product gas A ^* ) (caused by direct and / or indirect cooling).

Regardless of whether condensation removal of water from the product gas A is carried out, it is preferred that at least the propane circulation gas (and, if appropriate, the other components of the reaction gas A packed gas mixture Usually dehydrogenated recycle gas) is excluded), so that it is appropriate to heat the propane recycle gas (and, if appropriate, the other components of the reaction gas A packed gas mixture).

Alternatively, to use product gas A or product gas A ^* to fill one or more oxidation reactors in reaction zone B, an amount of oxygen molecules required to achieve the purpose in reaction zone B may be introduced into product gas A or product gas A ^{* Is} sufficient.

This addition is principally carried out by mixing pure oxygen or oxygen molecules with a mixture of one or more gases (e.g. N ₂ , H ₂ O, rare gas, CO ₂ ) which chemically inertly behaves in reaction zone B ) ≪ / RTI > (also referred to herein as the primary oxygen source). Preferably, according to the present invention, 30 vol% or less, preferably 25 vol% or less, advantageously 20 vol% or less, particularly advantageously 15 vol% or less, preferably 10 vol% or less, , More preferably no more than 5 vol%, or no more than 2 vol%. It is particularly advantageous to use pure oxygen at this time.

Typically, the amount of oxygen molecules added is such that the molar ratio of oxygen molecules to propylene present in the charge gas for reaction zone B (reaction gas B starting mixture) is greater than or equal to 1 and less than or equal to 3. Before adding the oxygen molecule-containing gas to the product gas A or the product gas A ^* , the temperature was advantageously adjusted to a value of 250 to 370 ° C, particularly advantageously 270 to 320 ° C. The gas containing the oxygen molecules is usually present in the supply to the reaction gas A or A ^* compressed to a pressure exceeding the pressure of the reaction gas A or the reaction gas A ^* (preferably 1 bar or less), and a gas containing oxygen molecules reaction to flow into the gas a or a ^* to be suppressed in a simple manner (the case of the oxygen source used is air, for example, typically compressed by the radial compressor, as described in DE-a 103 53 014). The temperature of the gas containing oxygen molecules is usually the temperature at which the desired temperature (often 240 to 300 ° C) is set for the reaction mixture B charged mixture. Generally, the temperature of this oxygen molecule-containing gas is 100 to 200 占 폚.

A typical reaction gas B charge gas mixture is

10 to 40% by volume of propane,

5 to 25% by volume of propylene,

0 to 10% by volume of H ₂ ,

10 to 30% by volume of O ₂ , and

1 to 10 and comprises a volume percentage of H ₂ O.

A preferred Reaction Gas B fill gas mixture comprises

15 to 40% by volume of propane,

7 to 20% by volume of propylene,

0 to 6% by volume of H ₂ ,

15 to 30% by volume of O ₂ , and

1 to 7 and comprises a H ₂ O volume%.

In the process according to the invention, the inlet pressure of the reaction gas B introduced into the at least one oxidation reactor of reaction zone B is usually suitably in the range from 1 bar to 4 bar, usually 1.3 bar to 3 bar, in most cases 1.5 bar 2.5 bar.

In a manner known per se, inhomogeneously catalyzed gaseous partial oxidation of propylene to acrylic acid using oxygen molecules is, in principle, carried out in accordance with the reaction scheme 2 in which the first stage induces acrolein and the second stage leads to acrylic acid from acrolein - Continue in steps.

This two-step sequential reaction sequence may be carried out in a manner known per se by terminating the process according to the invention in the acrolein step (predominant acrolein formation step) in reaction zone B and performing the target product removal in this step, To continue the process according to the invention and only open the possibility of carrying out the target product removal thereafter.

It is advantageous if the process according to the invention is carried out up to predominantly acrylic acid formation, in accordance with the invention, the process is carried out in a two-stage, i.e. continuously arranged, oxidation step, in which case a fixed catalyst bed to be used and preferably also Other reaction conditions, for example the temperature of the fixed catalyst bed, are suitably adjusted in an optimized manner for each of the oxidation stages.

A multimetal oxide comprising elements Mo, Fe and Bi, which is particularly suitable as the active composition for the first oxidation step (propylene to acrolein) catalyst, can also catalyze a second oxidation step (acrolein to acrylic acid) to some extent, A catalyst in which the composition is at least one multimetal oxide comprising elements Mo and V is preferred for the second oxidation step.

The method in which the catalyst is carried out in reaction zone B according to the invention for heterogeneously catalyzed partial oxidation of propylene on a fixed catalyst bed having at least one multimetal oxide containing elements Mo, Fe and Bi as the active composition, Step process for preparing acrylic acid, if appropriate) or as a first reaction step in the two-step preparation of acrylic acid.

The realization of one-step heterogeneously catalyzed partial oxidation of propylene to acrolein and acrylic acid, if appropriate, or two-step heterogeneously catalyzed partial oxidation of propylene to acrylic acid using reaction gas B obtained according to the present invention, EP-A 70 07 14 (the first reaction stage, also described therein, but also carried out in the corresponding countercurrent flow mode of the salt bath and the starting reaction gas mixture in the vortex reactor), EP-A 70 08 93 WO 04/085 369 (especially as this is considered as an integral part of this document) (as a two-step process), WO 04/85 363, DE- A 103 13 212 (first reaction step), EP-A 1 159 248 (as a two-step process), EP-A 1 159 246 (second reaction step) DE-A 199 48 248 (as a two-step process), DE-A 101 01 695 (one-step or two-step), WO 04/085 368 (As a two-step process), DE 10 2004 021 764 (two-step), WO 04/085362 (first reaction step), WO 04/085370 (second reaction step), WO 04/085365 ), WO 04/085367 (2-step), EP-A 990 636, EP-A 1 007 007 and EP-A 1 106 598.

This applies in particular to all embodiments included in the document. These can be carried out as described in the literature, but differ in that the starting reaction gas mixture used in the first reaction step (from propylene to acrolein) is the reaction gas B produced according to the invention. Regarding the remaining parameters (particularly with respect to the reactant loadings of the fixed catalyst bed and the fixed catalyst bed), the procedure is the same as in the example of the mentioned literature. When the procedure in the above-mentioned prior art example is in two stages and there is a secondary oxygen supply (secondary air supply in a less preferred manner in accordance with the invention) between the two reaction stages, the supply is carried out in an appropriate manner, The amount of oxygen molecules to acrolein in the packed gas mixture of the second reaction step is adjusted so as to correspond to the molar ratio of the embodiments of the mentioned reference.

Advantageously, according to the invention, the amount of oxygen in reaction zone B is such that product gas B still contains non-converting oxygen molecules (suitably from 0.5 to 6 vol%, advantageously from 1 to 5 vol% 2 to 4% by volume). For a two-step procedure, this applies to each of the two oxidation steps.

Multimetal oxide catalysts particularly suited for particular oxidation steps have been described several times in the past and are well known to those skilled in the art. For example, EP-A 253 409 represents a US patent corresponding to page 5.

Preferred catalysts for certain oxidation steps are also disclosed in DE-A 44 31 957, DE-A 10 2004 025 445 and DE-A 44 31 949. This applies in particular to the catalysts of the formula (I) in the above two references. Particularly advantageous catalysts for certain oxidation steps are described in DE-A 103 25 488, DE-A 103 25 487, DE-A 103 53 954, DE-A 103 44 149, DE-A 103 51 269, DE-A 103 50 822.

In heterogeneously catalyzed catalytic gas phase partial oxidation of propylene to acrolein or acrylic acid or mixtures thereof, the useful multimetal oxide compositions are essentially all multimetal oxide compositions comprising Mo, Bi and Fe as the active composition to be.

These include in particular the multimetal oxide active composition of formula I of DE-A 199 55 176, the multimetal oxide active composition of formula I of DE-A 199 48 523, the multimetal oxide active composition of formula I, II and III of DE-A 101 01 695 An oxide active composition, a multimetal oxide active composition of formulas I, II and III of DE-A 199 48 248 and a multimetal oxide active composition of formulas I, II and III of DE-A 199 55 168, 700 714. < / RTI >

Further, initiation Document No. 497 012 Lake (August 29, 2005) (inadvertently, in its embodiment, the specific surface area while having the same value incorrectly written to the correct m ² / g instead of in cm ² / g), DE A 100 46 957, DE-A 100 63 162, DE-C 3 338 380, DE-A 199 02 562, EP-A 15 565, DE-C 2 380 765, EP- 2793 74, DE-A 330 00 44, EP-A 575897, US-A 4438217, DE-A 19855913, WO 98/24746, DE-A 197 46 210 (catalyst of formula II), JP-A 91/294239 , EP-A 293 224, and Mo-Bi and Fe disclosed in EP-A 700 714 are suitable for this reaction step. This applies in particular to the exemplary embodiments of the document, of which catalysts of EP-A 15 565, EP-A 575 897, DE-A 197 46 210 and DE-A 198 55 913 are particularly preferred. In this regard, it has been found that catalysts prepared according to Example 1c of EP-A 15 565 and also correspondingly, but having an active composition of the composition Mo ₁₂ Ni _6.5 Zn ₂ Fe ₂ Bi ₁ P _0.0065 K _0.06 O _{x .10} SiO ₂ , do. In addition, as an unsupported hollow cylindrical catalyst having dimensions of 5 mm x 3 mm x 2 mm (outer diameter x height x inner diameter), an example of serial number 3 from DE-A 198 55 913 (stoichiometric formula: Mo ₁₂ Co ₇ Fe ₃ Bi _0.6 K _0.08 Si _1.6 O _x ) and also an unsupported multimetal oxide II catalyst according to Example 1 of DE-A 197 46 210. It is also necessary to mention the multimetal oxide catalysts of US-A 4,438,217. In the latter case, the dimensions of these hollow cylinders are 5.5 mm x 3 mm x 3.5 mm, or 5 mm x 2 mm x 2 mm, or 5 mm x 3 mm x 2 mm, or 6 mm x 3 mm x 3 mm, or 7 mm x 3 mm x 4 mm (each of outer diameter x height x inner diameter). Additional possible catalyst configurations in this regard are extrudates (e.g., 7.7 mm in length and 7 mm in diameter; or 6.4 mm in length and 5.7 mm in diameter).

A number of multimetal oxide active compositions suitable for the step from propylene to acrolein and, if appropriate, to acrylic acid, may be included in the following formula (IV).

Mo ₁₂ Bi _a Fe _b X ¹ _c X ² _d X ³ _e X ⁴ _f O _n

The variables are defined as follows:

X ¹ = nickel and / or cobalt,

X ² = thallium, an alkali metal and / or an alkaline earth metal,

X ³ = zinc, phosphorus, arsenic, boron, antimony, tin, cerium, lead and / or tungsten,

X ⁴ = silicon, aluminum, titanium and / or zirconium,

a = 0.5 to 5,

b = 0.01 to 5, preferably 2 to 4,

c = 0 to 10, preferably 3 to 10,

d = 0 to 2, preferably 0.02 to 2,

e = 0 to 8, preferably 0 to 5,

f = 0 to 10 and

n = number determined by valency and frequency of elements other than oxygen in formula IV.

They can be obtained in a manner known per se (see, for example, DE-A 4 023 239) and are usually shaped to provide substantially spheres, rings or cylinders, or in the form of coated catalysts, Lt; RTI ID = 0.0 > inert < / RTI > These may, of course, also be used as catalysts in powder form (for example in fluid bed reactors).

In principle, the active composition of formula (IV) can be prepared in a simple manner by obtaining a very tight, preferably finely divided, dry mixture having a composition corresponding to their stoichiometric formula from a suitable source of these element components and calcining them at a temperature of from 350 to 650 ° C. ≪ / RTI > Calcination may be carried out under an inert gas or an oxidizing atmosphere, for example air (a mixture of inert gas and oxygen) and also in a reducing atmosphere (e.g. a mixture of inert gas, NH ₃ , CO and / or H ₂ ). The calcination time can be from minutes to hours, and typically decreases with temperature. A useful source for the elemental constituents of the multi-metal oxide active composition IV is a compound which is already an oxide and / or a compound which can be converted into an oxide by heating in the presence of at least oxygen.

In addition to the oxides, such useful starting compounds include in particular halides, nitrates, formates, oxalates, citrates, acetates, carbonates, amine complexes, ammonium salts and / or hydroxides (which are decomposed and / NH ₄ OH, (NH ₄ ) ₂ CO ₃ , NH ₄ NO ₃ , NH ₄ CHO ₂ , CH ₃ COOH, NH ₄ CH ₃ CO ₂ and / or ammonium Compounds such as oxalate may be further incorporated into the intimate dry mix).

The starting compounds for preparing the multimetal oxide active composition IV may be intimately mixed in a dry or wet form. When they are mixed in dry form, the starting compounds are suitably used as finely divided powders and are calcined after mixing and optionally compressing. However, it is preferable to intimately mix in wet form. Typically, the starting compounds are mixed together in the form of an aqueous solution and / or an aqueous suspension. A particularly intimate dry mixture is obtained in the described mixing method when the starting material is the source of the elemental constituents in the exclusively dissolved form. The solvent used is preferably water. Thereafter, the obtained aqueous composition is dried, and the drying step is preferably carried out by spray-drying the aqueous mixture at an outlet temperature of 100 to 150 ° C.

The multimetal oxide active composition of formula (IV) can be formed in powder form or in the form of a particular catalyst and used in the "propylene to acrolein (and if appropriate acrylic acid)" step, and the molding can be carried out before or after the final calcination. For example, from the powder form of the active composition or its uncalcined and / or partially calcined precursor composition (for example by tableting or extrusion), as appropriate, auxiliary agents such as graphite or stearic acid as lubricant and / Unsupported catalysts can be prepared by adding molding aids and reinforcing agents, for example microfibers of glass, asbestos, silicon carbide or titanium potassium, to the desired catalyst geometry. Instead of graphite, it is also possible to use hexavalent boron nitride as auxiliary in molding, as recommended in DE-A 10 2005 037 678. Examples of suitable unsupported catalyst shapes include solid cylinders or hollow cylinders having an outer diameter and a length of 2 to 10 mm. For hollow cylinders, a wall thickness of 1 to 3 mm is advantageous. The unsupported catalyst may, of course, be spherical, and the spherical diameter may be between 2 and 10 mm. A particularly preferred hollow cylindrical shape is 5 mm x 3 mm x 2 mm (outer diameter x length x inner diameter), particularly for unsupported catalysts.

The powdered active composition or powder precursor composition that has already been calcined and / or partially calcined may, of course, be formed by applying it to a preformed inert catalyst support. The coating of the support on which the coated catalyst is prepared is generally carried out in a suitable rotatable vessel, for example as disclosed in DE-A 29 09 671, EP-A 293 859 or EP-A 714 700. To coat the support, the applied powder composition is suitably wetted and re-dried after application, e.g., with hot air. The coating thickness of the powder composition applied to the support is appropriately selected within the range of 10 to 1000 占 퐉, preferably 50 to 500 占 퐉, more preferably 150 to 250 占 퐉.

Useful support materials include conventional porous or nonporous aluminum oxide, silicon dioxide, thorium dioxide, zirconium dioxide, silicon carbide or silicates, such as magnesium silicate or aluminum silicate. They generally behave substantially inert to the target reaction on which the process according to the invention is based. A shaped support, such as a spherical or hollow cylinder, with significant surface roughness is preferred, but the support may be regular or irregular in shape. Suitable supports are spherical supports made of soap and having a substantially non-porous surface roughness of 1 to 10 mm or 8 mm, preferably 4 to 5 mm in diameter. However, a suitable support is also a cylinder having a length of 2 to 10 mm and an outer diameter of 4 to 10 mm. For suitable rings according to the invention as support, the wall thickness is also typically from 1 to 4 mm. The annular support preferably used according to the present invention has a length of 2 to 6 mm, an outer diameter of 4 to 8 mm and a wall thickness of 1 to 2 mm. A suitable support according to the invention is a ring of dimensions of 7 mm x 3 mm x 4 mm (outer diameter x length x inner diameter). The fineness of the catalytically active oxide composition to be applied to the surface of the support is of course adjusted according to the desired coating thickness (cf. EP-A 714 700).

The multimetal oxide active composition to be used in the step of propylene to acrolein (and, if appropriate, acrylic acid) is also a composition of formula V:

[Y ¹ _{a '} Y ² _b' O _{x '} ] _p [Y ³ _c' Y ⁴ _{d '} Y ⁵ _e' Y ⁶ _{f '} Y ⁷ _g' Y ² _{h '} O _y' ] _q

The variables are defined as follows:

(Y ¹ = only bismuth or bismuth and at least one of tellurium, antimony, tin and copper elements,

Y ² = molybdenum, or tungsten, or molybdenum and tungsten,

Y ³ = alkali metal, thallium and / or samarium,

Y ⁴ = alkaline earth metal, nickel, cobalt, copper, manganese, zinc, tin, cadmium and /

Y ⁵ = at least one of iron or iron and chromium and cerium elements,

Y ⁶ = phosphorus, arsenic, boron and / or antimony,

Y ⁷ = rare earth metal, titanium, zirconium, niobium, tantalum, rhenium, ruthenium, rhodium, silver, gold, aluminum, gallium, indium, silicon, germanium, lead, thorium and /

a '= 0.01 to 8,

b '= 0.1 to 30,

c '= 0 to 4,

d '= 0 to 20,

e '= 0 to 20,

f '= 0 to 6,

g '= 0 to 15,

h '= 8 to 16,

x ', y' = number determined by valence and frequency of elements other than oxygen in formula V, and

p, q = p / q ratio of 0.1 to 10)

(The longest straight line connecting the two points on the surface (interface) of the region passing through the center of the region) of 1 nm to 100 μm, often 10 nm to 500 nm, or a three-dimensional region of chemical composition Y ¹ _{a '} Y ² _b' O _{x '} of 1 μm to 50 or 25 μm.

A particularly advantageous multimetal oxide composition V according to the invention is a composition wherein Y < ^{1 >} is simply bismuth.

Among them, a composition represented by the following formula (VI) is preferred.

^{_{[Bi a "Z 2 b"}} O x "] p" [Z 2 12 Z 3 c "Z 4 d" Fe e "Z 5 f" Z 6 g "Z 7 h" O y "] q"

The variables are defined as follows:

Z ² = molybdenum, or tungsten, or molybdenum and tungsten,

Z ³ = nickel and / or cobalt,

Z ⁴ = thallium, an alkali metal and / or an alkaline earth metal,

Z ⁵ = phosphorus, arsenic, boron, antimony, tin, cerium and / or lead,

Z ⁶ = silicon, aluminum, titanium and / or zirconium,

Z ⁷ = Copper, silver and / or gold,

a "= 0.1 to 1,

b "= 0.2 to 2,

c "= 3 to 10,

d "= 0.02 to 2,

e "= 0.01 to 5, preferably 0.1 to 3,

f "= 0 to 5,

g "= 0 to 10,

h "= 0 to 1,

x ", y "= number determined by the valence and frequency of elements other than oxygen in formula (VI)

quot; p ", q "= p" / q "is 0.1 to 5, preferably 0.5 to 2.

Very particularly preferred is composition VI wherein Z ² _{b "} = (tungsten) _b" and Z ² ₁₂ = (molybdenum) ₁₂ .

[Y ¹ _{a '} Y ² _b' O _{x '} ) of a suitable multimetal oxide composition V (multimetal oxide composition VI) according to the present invention in a suitable multimetal oxide composition V (multimetal oxide composition VI) (preferably 50 mol% or more, and more preferably 100 mol% or more) of the total proportion of [ _p ] ([Bi _{a "} Z ² _b" O _{x "} ] _p" It is in the form of a three-dimensional region of the chemical composition Y ¹ _{a '} Y ² _b' O _{x '} [Bi _{a "} Z ² _b" O _{x "} ] which has a boundary with its local environment due to its chemical composition, To 100 [mu] m.

In molding, the reference to the multimetal oxide composition IV catalyst is applied to the multimetal oxide composition V catalyst.

The preparation of the multimetal oxide active composition V is described, for example, in EP-A 575 897 and also in DE-A 198 55 913.

The recommended inert support material is also particularly useful as an inert material for the dilution and / or boundary of a suitable fixed catalyst bed, or as an upstream layer protecting them and / or heating the gas mixture.

In the case of the second step (second reaction step), the active composition useful for catalysts required for heterogeneously catalyzed gaseous partial oxidation of acrolein to acrylic acid is, as described above, in principle, all multimetal oxide compositions comprising Mo and V , For example DE-A 100 46 928.

A number of compositions, for example a composition of DE-A 198 15 281, may be included in formula VII below.

Mo ₁₂ V _a X ¹ _b X ² _c X ³ _d X ⁴ _e X ⁵ _f X ⁶ _g O _n

The variables are defined as follows:

X ¹ = W, Nb, Ta, Cr and / or Ce,

X ² = Cu, Ni, Co, Fe, Mn and / or Zn,

X ³ = Sb and / or Bi,

X ⁴ = at least one alkali metal,

X ⁵ = at least one alkaline earth metal,

X ⁶ = Si, Al, Ti and / or Zr,

a = 1 to 6,

b = 0.2 to 4,

c = 0.5 to 18,

d = 0 to 40,

e = 0 to 2,

f = 0 to 4,

g = 0 to 40, and

n = number determined by valency and frequency of elements other than oxygen in formula (VII).

Preferred embodiments of the active polymetal oxide VII according to the present invention are those wherein the variables of formula VII are included in the definition:

X ¹ = W, Nb and / or Cr,

X ² = Cu, Ni, Co and / or Fe,

X ³ = Sb,

X ⁴ = Na and / or K,

X ⁵ = Ca, Sr and / or Ba,

X ⁶ = Si, Al and / or Ti,

a = 1.5 to 5,

b = 0.5 to 2,

c = 0.5 to 3,

d = 0 to 2,

e = 0 to 0.2,

f = 0 to 1, and

n = number determined by valency and frequency of elements other than oxygen in formula (VII).

However, a very particularly preferred multimetal oxide VII according to the present invention is an oxide of the following formula VIII.

Mo ₁₂ V _{a '} Y ¹ _b' Y ² _{c '} Y ⁵ _f' Y ⁶ _{g'O n '}

In this formula,

Y ¹ = W and / or Nb,

Y ² = Cu and / or Ni,

Y ⁵ = Ca and / or Sr,

Y ⁶ = Si and / or Al,

a '= 2 to 4,

b '= 1 to 1.5,

c '= 1 to 3,

f '= 0 to 0.5

g '= 0 to 8, and

n '= number determined by valency and frequency of elements other than oxygen in formula VIII.

A suitable multimetal oxide active composition VII according to the invention can be obtained in a manner known per se, for example as disclosed in DE-A 43 35 973 or EP-A 714 700.

In principle, a multimetal oxide active composition suitable for the "acrolein to acrylic acid" step, in particular a composition of formula VII, can be prepared from a suitable source of their elemental components by a very intimate, preferably finely divided, dry mixture having a composition corresponding to their stoichiometry And calcining it at a temperature of from 350 to 600 < 0 > C. The calcination may be carried out under an inert gas or an oxidizing atmosphere, for example air (mixture of inert gas and oxygen) and also in a reducing atmosphere (for example inert gas and reducing gas such as H ₂ , NH ₃ , CO, methane and / Acrolein or a mixture of reducing gases as such). The calcination time can be from several minutes to several hours, and typically decreases with temperature. A useful source for the elemental constituents of the multimetal oxide active composition VII comprises compounds which are already oxides and / or compounds which can be converted into oxides by heating in the presence of at least oxygen.

The starting compounds for the preparation of the multimetal oxide composition VII can be intimately mixed in dry or wet form. When they are mixed in dry form, the starting compounds are suitably used in the form of finely divided powders, mixed and, where appropriate, calcined after compression. However, it is preferable to intimately mix in wet form.

This is typically done by mixing the starting compounds together in the form of aqueous solutions and / or aqueous suspensions. A particularly intimate dry mixture is obtained by the described mixing method when the starting material is exclusively the source of the elemental constituents in dissolved form. The solvent used is preferably water. Thereafter, the obtained aqueous composition is dried, and the drying step is preferably carried out by spray-drying the aqueous mixture at an outlet temperature of 100 to 150 ° C.

The resulting multimetal oxide composition, in particular the composition of formula VII, can be used for acrolein oxidation in powder form (e.g. in a fluidized bed reactor) or in a form that is molded into a specific catalyst form, and the molding can be performed before or after the final calcination . For example, an unsupported catalyst can be prepared from the powder form of the active composition or its uncalcined precursor composition (e.g., by tableting or extrusion), if appropriate, with adjuvants such as graphite or stearic acid and / The unsupported catalyst can be prepared by adding a reinforcing agent, for example, fine fibers of glass, asbestos, silicon carbide or titanium potassium to the desired catalyst shape. Examples of suitable non-supported catalyst shapes are solid cylinders or hollow cylinders having an outer diameter and a length of 2 to 10 mm. In the case of a hollow cylinder, a wall thickness of 1 to 3 mm is suitable. The unsupported catalyst may, of course, have a spherical shape, in which case the sphere diameter may be between 2 and 10 mm (e.g. 8.2 or 5.1 mm).

The powder active composition or powder precursor composition to be calcined may, of course, be formed by applying to a preformed inert catalyst support. The coating of the support on which the coated catalyst is prepared is generally carried out in a suitable rotatable vessel, for example as disclosed in DE-A 2 909 671, EP-A 293 859 or EP-A 714 700.

To coat the support, the powder composition to be applied is suitably wetted and re-dried, for example, with hot air, after application. The coating thickness of the powder composition applied to the support is appropriately selected within the range of 10 to 1000 占 퐉, preferably 50 to 500 占 퐉, more preferably 150 to 250 占 퐉.

Useful support materials include conventional porous or nonporous aluminum oxide, silicon dioxide, thorium dioxide, zirconium dioxide, silicon carbide or silicates, such as magnesium silicate or aluminum silicate. A flat support having a significant surface roughness, for example a sphere or hollow cylinder with a grit layer, is preferred, but the support may be regular or irregular in shape. Suitable supports include substantially non-porous roughened spherical supports made of soap and having a diameter of 1 to 10 mm or 8 mm, preferably 4 to 5 mm. That is, the suitable spheres may be 8.2 mm or 5.1 mm in diameter. However, suitable supports also include a cylinder having a length of 2 to 10 mm and an outer diameter of 4 to 10 mm. In the case of a ring as a support, the wall thickness is also typically from 1 to 4 mm. The annular support which is preferably used has a length of 2 to 6 mm, an outer diameter of 4 to 8 mm and a wall thickness of 1 to 2 mm. A suitable support is also a ring of dimensions of 7 mm x 3 mm x 4 mm (outer diameter x length x inner diameter). The fineness of the catalytically active oxide composition to be applied to the surface of the support is of course adjusted according to the desired coating thickness (cf. EP-A 714 700).

A preferred multimetal oxide active composition for use in the "acrolein to acrylic acid" step is also a composition of formula IX:

It is also possible to preform separately the preformed solid starting composition 1 into the following polymorphic form D of the following polymetallic oxide composition E (starting composition 1) in the form of fine particles, then add the elements Mo, V, Z ¹ , Z ² , Z ³ , In an aqueous solution, aqueous suspension or finely divided dry mixture (starting composition 2) of said elemental source comprising Z ⁴ , Z ⁵ , Z ⁶ at the desired p: q ratio, drying the resulting aqueous mixture, And calcining the resulting dried precursor composition before or after drying at a temperature of 250 to 600 DEG C to obtain the desired catalyst shape.

[D] _p [E] _q

The variables are defined as follows:

D = Mo ₁₂ V _{a "} Z ¹ _b" Z ² _{c "} Z ³ _d" Z ⁴ _{e "} Z ⁵ _f" Z ⁶ _{g "} O _x"

E = Z ⁷ ₁₂ Cu _{h "} H _i" O _{y "} ,

Z ¹ = W, Nb, Ta, Cr and / or Ce,

Z ² = Cu, Ni, Co, Fe, Mn and / or Zn,

Z ³ = Sb and / or Bi,

Z ⁴ = Li, Na, K, Rb, Cs and / or H,

Z ⁵ = Mg, Ca, Sr and / or Ba,

Z ⁶ = Si, Al, Ti and / or Zr,

Z ⁷ = Mo, W, V, Nb and / or Ta, preferably Mo and / or W,

a "= 1 to 8,

b "= 0.2 to 5,

c "= 0 to 23,

d "= 0 to 50,

e "= 0 to 2,

f "= 0 to 5,

g "= 0 to 50,

h "= 4 to 30,

i "= 0 to 20,

x ", y "= number determined by valency and frequency of elements other than oxygen in formula (IX), and

p, q = p / q ratio of 160: 1 to 1: 1.

Z ⁷ ₁₂ Cu _{H "} H _i" O _{y "}

Mo ₁₂ V _{a "} Z ¹ _b" Z ² _{c "} Z ³ _d" Z ⁴ _{e "} Z ⁵ _f" Z ⁶ _{g "}

Preference is given to a polymetallic oxide composition IX in which the preformed solid starting composition 1 is incorporated into the aqueous starting composition 2 at a temperature less than < RTI ID = 0.0 > 70 C. < / RTI & A detailed description of the preparation of multicomponent oxide composition VI catalysts is contained, for example, in EP-A 668 104, DE-A 197 36 105, DE-A 100 46 928, DE-A 197 40 493 and DE-A 195 28 646 have.

In molding, a reference to a multimetal oxide composition VII catalyst is applied to the multimetal oxide composition IX catalyst.

Multimetal oxide catalysts which are very suitable for the "acrolein to acrylic acid" stage are also catalysts of DE-A 198 15 281, in particular those having a multimetal oxide active composition of the formula I herein.

Advantageously, an unsupported catalyst ring is used in the step from propylene to acrolein, and the coated catalyst ring is used in the step from acrolein to acrylic acid.

The performance of the partial oxidation process according to the invention from propylene to acrolein (and, if appropriate, acrylic acid) is carried out using the catalysts described in the 1-zone multi-catalyst tube fixed bed reactor as described for example in DE-A 44 31 957 . In such a case, the reaction gas mixture and the heat carrier (heat exchange medium) can be passed as co-current or counter current when observed on the reactor.

The reaction pressure is typically 1 to 3 bar and the total space velocity on the fixed catalyst bed of reaction gas B is preferably 1500 to 4000 or 6000 l (STP) / l · h or more. The propylene loading (propylene time-space velocity on the fixed catalyst bed) is typically between 90 and 200 l (STP) / l · h or 300 l (STP) / l · h or more. Propylene loading greater than or equal to 135 l (STP) / l · h or 140 l (STP) / l · h or 150 l (STP) / l · h or 160 l (STP) Since the starting reaction gas mixture for reaction zone B of the present invention results in the desired hot spot behavior due to the presence of unconverted propane and, where appropriate, hydrogen molecules, all of which are also applied irrespective of the particular choice of fixed bed reactor, .

The flow of the packed gas mixture into the 1-zone multi-catalyst tube fixed bed reactor is preferably from above. The heat exchange medium used is suitably a salt melt, preferably consisting of 60 wt.% Of potassium nitrate (KNO ₃ ) and 40 wt.% Of sodium nitrite (NaNO ₂ ) or 53 wt.% Of potassium nitrate (KNO ₃ ) 40% by weight of sodium (NaNO ₂ ) and 7% by weight of sodium nitrate (NaNO ₃ ).

Upon observation on the reactor, the salt melt and the reaction gas mixture may be delivered as co-current or countercurrent as described above. The salt melt itself is preferably delivered in a meandering manner around the catalyst tube.

When the flow from the top to the bottom of the catalyst tube is low, it is appropriate to charge the catalyst in the catalyst tube from bottom to top as follows (in the case of flow from bottom to top, the filling order is reversed appropriately).

First, a catalyst or a mixture of a catalyst and an inert material (the inert material constitutes a weight ratio of 30 or 20% by weight or less, based on the mixture) to a length of 40 to 80 or 60% C);

Subsequently, a catalyst or a mixture of a catalyst and an inert material (the inert material constitutes a weight ratio of not more than 40% by weight based on the mixture) (segment B) to a length of 20 to 50 or 40% ; And

Finally, a layer of inert material (compartment A) which is chosen to cause a pressure drop of preferably 10 to 20% of the total tube length, preferably a very small pressure drop.

Zone C is preferably not diluted.

The above-mentioned charge modification method is characterized in that the catalyst used is a catalyst according to the method described in the publication No. 497012 (Aug. 29, 2005) or Example 1 of DE-A 100 46 957 or Example 3 of DE-A 100 46 957 , And the inert material used is particularly suitable for the case of a copper sheet having dimensions of 7 mm x 7 mm x 4 mm (outer diameter x height x inner diameter). As regards the salt bath temperature, the description of DE-A 4 431 957 applies.

However, the partial oxidation of propylene to acrolein (and, if appropriate, acrylic acid) in reaction zone B can also be carried out in the presence of a catalyst such as, for example, DE-A 199 10 506, DE-A 10 2005 009 885, DE-A 10 2004 032 129, DE- A 10 2005 013 039 and DE-A 10 2005 009 891, and also DE-A 10 2005 010 111, for example using a catalyst as described in a two-zone multi-catalyst tube fixed bed reactor. In all of the cases described above (and very generally in the process according to the invention), the propene conversion achieved in single pass is typically at least 90 mol%, or at least 95 mol%, and the acrolein formation selectivity is at least 90 mol% . Advantageously, according to the invention, the partial oxidation of propylene to acrolein or acrylic acid or mixtures thereof according to the invention is carried out as described in EP-A 1 159 244, most preferably in WO 04/085363 and WO 04/085362 .

The documents EP-A 1 159 244, WO 04/085363 and WO 04/085362 are regarded as an integral part of the present application.

In other words, the partial oxidation of propylene to acrolein (and, if appropriate, acrylic acid) to be carried out in reaction zone B can be carried out particularly advantageously on a fixed catalyst bed having an increased propylene loading and two or more temperature zones.

In this connection, reference is made, for example, to EP-A 1 159 244 and WO 04/085362.

The second stage of the two-stage partial oxidation of propylene to acrylic acid, i. E. The partial oxidation of acrolein to acrylic acid, is carried out in a 1-zone multi-catalyst tube fixed bed reactor as described for example in DE-A 44 31 949 . In this reaction step, the reaction gas mixture and the thermal carrier can be delivered as co-currents when observed on the reactor. Generally, the product gas mixture of the propylene partial oxidation of the preceding propylene partial oxidation of the invention to the acrolein is obtained in principle by itself (if appropriate after intermediate cooling, for example by addition of secondary oxygen (or, in a less preferred manner, (I.e., indirectly or directly by the addition of air), i.e., without secondary component removal, into the second reaction stage, i.e., acrolein partial oxidation.

The second step, the oxygen molecule required for acrolein partial oxidation, may already be present in the reaction gas B starting mixture for the partial oxidation of propylene to acrolein. However, it can also be added partially or completely directly to the first reaction step, i.e. to the product gas mixture of the partial oxidation of propylene to acrolein (which can be carried out in the form of secondary air, preferably pure oxygen or inert (Preferably 50 vol% or less, or 40 vol% or less, or 30 vol% or less, or 20 vol% or less, or 10 vol% or less, or 5 vol% or less , Or 2% by volume or less). In the case of pressure and temperature of the second oxygen source, the same applies as described for the first oxygen source. Regardless of the procedure, the packed gas mixture (starting reaction gas mixture) of acrolein partial oxidation to this acrylic acid advantageously has the following contents:

3 to 25% by volume of acrolein,

5 to 65% by volume of oxygen molecules,

6 to 70% by volume of propane,

0 to 20% by volume of a hydrogen molecule,

8 to 65% by volume of water vapor, and

0 to 70% by volume nitrogen molecule.

The starting reaction gas mixture preferably has the following contents:

5 to 20% by volume of acrolein,

5 to 15% by volume of oxygen molecules,

10 to 40% by volume of propane,

0 to 10% by volume of a hydrogen molecule, and

Water vapor 5 to 30% by volume.

The starting reaction gas mixture most preferably has the following contents:

10 to 20% by volume of acrolein,

7 to 15% by volume of oxygen molecules,

20 to 40% by volume of propane,

3 to 10% by volume of a hydrogen molecule, and

Water vapor 15 to 30% by volume.

The nitrogen content in the mixture will generally be no greater than 20 vol.%, Preferably no greater than 15 vol.%, More preferably no greater than 10 vol.%, Most preferably no greater than 5 vol.%. The molar ratio of O ₂ to acrolein (O ₂ : acrolein) present in the packed gas mixture for the second oxidation step is advantageously generally greater than or equal to 0.5 and less than or equal to 2, often greater than or equal to 1 and less than or equal to 1.5 according to the present invention.

The CO ₂ content of the packed gas mixture for the second oxidation step will generally be no more than 5% by volume.

The reaction pressure in the second reaction step (acrolein to acrylic acid) is also typically in the range from 1 to 4 bar, preferably 1.3 to 3 bar, in many cases 1.5 to 2.5 bar, as in the first reaction step (propylene to acrolein) And the total space velocity on the fixed catalyst bed of the (starting) reaction gas mixture is preferably 1500 to 4000 or 6000 l (STP) / l · h or more. The acrolein loading (acrolein time-space velocity on the fixed catalyst bed) is typically between 90 and 190 l (STP) / l · h or 290 l (STP) / l · h or higher. The amount of acrolein loaded above 135 l (STP) / l · h or 140 l (STP) / l · h or above 150 l (STP) / l · h or above 160 l (STP) Particularly preferred is the starting reaction gas mixture to be used in accordance with the invention, since it results in similarly favorable hot spot behavior due to the presence of propane and, where appropriate, hydrogen molecules.

The temperature of the packed gas mixture for the second oxidation step is generally from 220 to 270 캜.

The conversion of acrolein is suitably usually at least 90 mol% based on the single pass of the starting reaction gas mixture through the fixed catalyst layer in the second oxidation step, and the selectivity of formation of acrylic acid to be accompanied is at least 90 mol%.

The flow of the packed gas mixture into the 1-zone multi-catalyst tube fixed bed reactor is likewise preferably from the top. The heat exchange medium used in the second step is also suitably a salt melt, preferably 60 wt.% Potassium nitrate (KNO ₃ ) and 40 wt.% Sodium nitrite (NaNO ₂ ) or 53 wt.% Potassium nitrate (KNO ₃ ) 40 wt% of sodium nitrite (NaNO ₂ ), and 7 wt% of sodium nitrate (NaNO ₃ ). As noted above, upon observation on the reactor, the salt melt and the reaction gas mixture may be delivered in cocurrent or countercurrent. The salt melt itself is preferably delivered in a mandrel fashion around the catalyst tube.

When the flow to the catalyst tube is low at the top, it is appropriate to fill the catalyst from bottom to top as follows:

First, the catalyst or a mixture of the catalyst and the inert material (the inert material constitutes a weight ratio of 30 (or 20) wt% or less based on the mixture) to a length of 50 to 80 or 70% Compartment C);

Then, in a length of 20 to 40% of the total tube length, only the catalyst or a mixture of catalyst and inert material (the inert material constitutes a proportion of 50 or 40% by weight or less) (compartment B); And

Finally, a layer of inert material (compartment A) which is chosen to cause a pressure drop of preferably 5 to 20% of the total tube length, preferably very small.

Zone C is preferably not diluted. In the case of heterogeneously catalyzed gaseous partial oxidation of acrolein to acrylic acid, particularly at high acrolein loading on the catalyst bed and high water vapor content of the packed gas mixture, Zone B can also consist of two successive catalyst diluents (To minimize hot spot temperature and hot spot temperature sensitivity). From the bottom to the top, first 30 (or 20)% by weight of inert material is filled, and then more than 20% by weight to 50 or 40% by weight of inert material. Subsequently, the compartment C is preferably not diluted.

In the case of flow from the bottom to the top into the catalytic tube, the catalytic tube charge is suitably reversed.

The above-mentioned fill regime is particularly suitable when the catalyst used is a catalyst according to Preparation Example 5 of DE-A 100 46 928 or a catalyst according to DE-A 198 15 281, and the inert material used has dimensions of 7 mm x 7 mm x 4 mm or 7 mm x 7 mm x 3 mm (in each case, outer diameter x height x inner diameter). As regards the salt bath temperature, the description of DE-A 443 19 49 applies. Is generally selected in such a way that the acrolein conversion achieved in a single pass is usually at least 90 mol%, or at least 95 mol%, or at least 99 mol%.

However, partial oxidation of acrolein to acrylic acid can also be carried out using the catalysts described, for example, in a two-zone multi-catalyst tube fixed bed reactor as described in DE-A 199 10 508. For acrolein conversion, the above description applies. Also, when partial oxidation of acrolein as described above is performed as the second reaction step of the two-step propylene partial oxidation to acroleic acid in the two-zone multi-catalyst tube fixed bed reactor, the packed gas mixture (starting reaction gas mixture) (Indirectly or indirectly) by suitably using a product gas mixture of partial oxidation (as already described above) guided to the first stage (propylene to acrolein) After cooling). The oxygen required for acrolein partial oxidation is preferably present in the form of pure oxygen molecules (if appropriate also in the form of air, or in the form of a mixture of oxygen molecules and an inert gas, preferably 50 vol% or less, 40 vol% or less, or 30 vol% or less, or 20 vol% or less, more preferably 10 vol% or less, or 5 vol% or less, or 2 vol% - stage The first stage of partial oxidation (propylene to acrolein) is added directly to the product gas mixture. However, it may already be present in the starting reaction gas mixture for the first reaction step as already described.

In a two-stage partial oxidation of propylene to acrylic acid using a further addition of the product gas mixture of the first stage of partial oxidation to fill the second stage of partial oxidation, two 1-zone multi-catalyst tube fixed bed reactors In the case of high reactant loading on the catalyst bed as in the case, the co-current mode between the reaction gas and the salt (thermal carrier) is preferred when viewed on the vortex reactor or two two-zone multicatalyst tube fixed bed reactors will generally be cascaded . Mixed series connections (1-zone / 2-zone or vice versa) are also possible.

Intermediate coolers, which may, where appropriate, include an inert layer capable of performing a filtering function, may be disposed between the reactors. The salt bath temperature of the multi-catalyst tube reactor for the first stage of the propylene two-step partial oxidation to acrylic acid is generally from 300 to 400 ° C. The salt bath temperature of the multi-catalyst tube reactor for acrolein partial oxidation to acrylic acid, which is the second step of the partial oxidation of propylene to acrylic acid, is usually from 200 to 350 ° C. In addition, the heat exchange medium (preferably a salt melt) is typically passed through the associated multi-catalyst tube fixed bed reactor in an amount such that the difference between the inlet temperature and the outlet temperature thereof is generally below 5 占 폚. As already mentioned, both steps of the partial oxidation of propylene to acrylic acid can be carried out in one reactor on one charge as described in DE-A 101 21 592.

It should also be mentioned here that some of the starting gas B starting mixture for the first stage ("propylene to acrolein") may be residual gas from separation zone II.

It should be emphasized that for the sake of perfection, the new propane can be metered additionally to either or both of the fill gas mixture for the first oxidation step and the fill gas mixture for the second oxidation step. This is undesirable for the present invention, but may be preferred in some cases, to exclude flammability of the packed gas mixture.

Overall, a tube bundle reactor in which the catalyst charge is appropriately changed along the individual catalyst tubes with the completion of the first reaction step (such a two-stage propylene partial oxidation in a single reactor is described, for example, in EP-A 911 313, EP-A 979 813, EP-A 990 636 and DE-A 28 30 765) form the simplest embodiment of two oxidation steps for two stages of partial oxidation of propylene to acrylic acid. If appropriate, a layer of inert material is interposed between the catalyst charge of the catalyst tube.

However, it is desirable to perform two oxidation steps in the form of two cascaded tube bundles systems. They can be arranged in one reactor and in this case the transition from one tube bundle to another is formed by a layer of inert material which is not received in the catalyst tube (suitably accessible at foot). The catalytic tube generally flows around by the heat carrier, while it does not reach the layer of inert material received as described above. Advantageously, therefore, the two catalyst bundles are accommodated in a spatially separated reactor. Generally, an intercooler is disposed between the two bundle reactors to reduce any post-acrolein post combustion progression in the product gas mixture leaving the first reaction zone. In the first reaction step (propylene → acrolein), the reaction temperature is generally 300 to 450 ° C, preferably 320 to 390 ° C. In the second reaction step (acrolein to acrylic acid), the reaction temperature is generally 200 to 370 ° C, often 220 to 330 ° C. In both oxidation steps, the reaction pressure is suitably from 1 to 4 bar, advantageously from 1.3 to 3 bar. The loading amount (l (STP) / l · h) of the reaction gas on the oxidation catalyst in both reaction stages is often 1500 to 2500 l (STP) / l · h or 4000 l (STP) / l · h. The loading amount of propylene or acrolein may be 100 to 200 or 300 l (STP) / l · h or more.

In principle, the two oxidation steps in the process according to the invention can be carried out, for example, as described in DE-A 198 37 517, DE-A 199 10 506, DE-A 199 10 508 and DE-A 198 37 519 .

In all reaction steps, excess oxygen molecules have a favorable effect on the reaction rate and the catalyst life of the specific gas phase partial oxidation compared to the amount required according to the reaction stoichiometry.

In principle, it is also possible to realize inhomogeneously catalyzed gaseous partial oxidation of propylene to acrylic acid to be carried out in accordance with the invention in a single zone vortex reactor as follows. Both reaction steps are carried out in an oxidation reactor filled with one or more catalysts in which the active composition is a multimetal oxide comprising elemental Mo, Fe and Bi and capable of catalyzing the reaction in both reaction steps. The catalyst charge can, of course, change continuously or suddenly along the reaction coordinates. Of course, in one embodiment of the two-stage partial oxidation of propylene to acrylic acid of the present invention in the form of two oxidation stages in series, the product formed as a by-product in the first oxidation step and, if necessary, It is possible to partially or completely remove the carbon dioxide and water vapor present in the product gas mixture exiting the first oxidation step from the gas mixture. It is desirable to select a procedure that does not provide such removal in accordance with the present invention.

A useful source for the intermediate oxygen supply carried out between the two oxidation stages is a pure oxygen molecule or CO ₂ , CO, noble gas, N ₂ and / or N ₂ , in addition to air (which is less preferred in accordance with the invention) Or an oxygen molecule diluted with an inert gas such as saturated hydrocarbons. Preferably, according to the invention, the oxygen source is at most 50 vol%, preferably at most 40 vol%, more preferably at most 30 vol%, even more preferably at most 20 vol%, preferably at most 10 vol% Or 5 vol% or less, or 2% or less of oxygen molecules.

The material useful for the inert diluted shaped body in the catalyst bed for the two oxidation steps is also a recommended material for the dilution of the catalyst bed in reaction zone A.

In the process according to the invention, for example, the metering of the cold oxygen source to the product gas mixture of the first partial oxidation stage can be carried out directly prior to being used as a constituent of the starting reaction gas mixture for the second partial oxidation step And the like.

Advantageously, according to the invention, the partial oxidation of acrolein to acrylic acid is carried out as described in EP-A 1 159 246, most preferably in WO 04/085365 and WO 04/085370. However, according to the present invention there is provided a starting reaction gas mixture comprising acrolein, the product gas mixture of propylene 1-step partial oxidation to acrolein of the present invention and, where appropriate, the ratio of oxygen molecules to acrolein in the resulting starting reaction gas mixture, It is preferred to use a starting reaction gas mixture supplemented with sufficient secondary air such that it is between 0.5 and 1.5. The documents EP-A 1 159 246, WO 04/08536 and WO 04/085370 are regarded as an integral part of this document.

In other words, the acrolein partial oxidation of the present invention to acrylic acid can be carried out on a fixed catalyst bed having advantageously two or more temperature zones with increased acrolein loading.

Overall, the propylene two-step partial oxidation to acrylic acid will preferably be carried out as described in EP-A 1 159 248 or WO 04/085367 or WO 04/085369.

The product gas B leaving the partial oxidation to be carried out according to the invention (after the first and / or second oxidation step) can be carried out in the presence of acrolein or acrylic acid or mixtures thereof as the target product, unconverted propane, if appropriate hydrogen molecules, (And / or further used as diluent gas), any non-converting oxygen molecules (in terms of the lifetime of the catalyst used, the oxygen content in the product gas mixture of all the partial oxidation stages is still, for example, still in the range from 1.5 to 4% And also other by-products or secondary components having a boiling point above or below water, such as CO, CO ₂ , lower aldehydes, lower alkanecarboxylic acids (such as acetic acid, formic acid and propionic acid ), Maleic anhydride, benzaldehyde, aromatic carboxylic acid and aromatic carboxylic acid anhydride (e.g., phthalic anhydride and benzoic acid) If adds hydrocarbons such as C ₄ hydrocarbons substantially free of a N ₂₎ diluted with an inert gas, for an additional (for example, the butene-1 and possibly other butenes), and in some cases with. A process useful for the removal of the target product present in the product gas B, generally accompanied by simultaneous removal of water and a second component of boiling point higher than water, for the process according to the invention is, in principle, All the way. Essential features of these processes include the conversion of the target product into a condensation phase, for example by absorption and / or condensation means, and subsequent extraction, distillation, crystallization and / or desorption of the condensate or absorber for further target product removal And then post-processed by the means. With the conversion of the target product to the target product and / or subsequent condensation phase, these processes typically also remove water vapor and the secondary component, which is a boiling point higher than water present in reaction gas B, into a condensed phase (preferably Is preferably at least 70 mol%, preferably at least 80 mol%, preferably at least 90 mol%, more preferably at least 90 mol% of the amount of water vapor formed in the reaction zone B (or more preferably the total amount present in the product gas B) More preferably at least 95 mol%, is converted to a condensed phase in the separation zone II).

Useful absorbents include, for example, water, aqueous solutions and / or organic solvents (for example mixtures of diphenyl and diphenyl ethers, or mixtures of diphenyl, diphenyl ether and o-dimethyl phthalate). The "condensation" (removal) of the secondary product having a higher boiling point than the target product, water and water, is usually carried out in the presence of a product gas B having relatively low boiling point and lower condensation, leaving residual gas that has not been converted into a condensation phase ≪ / RTI > These are typically components which have a boiling point of not more than -30 ° C, especially at standard pressure (1 bar) (their total content in the residual gas is generally at least 60% by volume, at least 70% by volume, in many cases at least 80% by volume, Usually 90% by volume or less). These unconverted propane, carbon dioxide, and any unconverted propylene, any unconverted oxygen molecules, and where appropriate, hydrogen molecules, and other secondary components having a lower boiling point than water, such as CO, ethane, methane, in some N ₂ , And rare gases (e.g., He, Ne, Ar, etc.) in some cases. The residual gas may also contain acrylic acid, acrolein and / or H ₂ O to some extent. Preferably, according to the invention, the residual gas comprises not more than 10% by volume, advantageously not more than 5% by volume, particularly advantageously not more than 3% by volume of water vapor. This residual gas is typically greater than (typically 80% or more, or 90% or more, or 95% or more, or more) of the residual gas formed in the second separation zone II (based on the amount of propane present therein) And is referred to herein as the (main) residual gas.

In particular, when the target product is condensed by absorption by one organic solvent, one or more secondary residual gases comprising non-converted propane and any unconverted propylene are generally obtained in the separation zone II (the propane (Generally not more than 20%, usually not more than 10%, or not more than 5%, or not more than 1%) as compared with the (main) residual gas amount. It is also due to the condensation phase which forms and absorbs a certain amount of unconverted propane and any unconverted propylene.

During the further course of extraction, distillation, crystallization and / or desorption removal (from separation zone II) of the target product from the condensation phase, the unconverted propane and, where appropriate, propylene are usually recovered as components of one or more additional gaseous phases, (Secondary) residual gas.

In this case, the sum of the (main) residual gas and the (secondary) residual gas forms the total amount of residual gas remaining in the separation zone II. (Secondary) residual gas is obtained in the separation zone II, the (main) residual gas is automatically the total amount of residual gas (also known as (total) residual gas).

Preferably, according to the present invention, the target product is converted by fractional condensation from the product gas B into a condensed phase. This is particularly the case when the target product is acrylic acid. However, suitable methods for the removal of the target products are, for example, those described in documents DE-A 102 13 998, DE-A 22 63 496, US 3,433,840 (the methods described above are particularly recommended for acrolein removal) EP-A 1 385 532, DE-A 102 35 847, EP-A 792 867, WO 98/01415, EP-A 1 015 411, EP-A 1 015 410, WO 99/50219, WO 00 / DE-A 102 43 625, DE-A 103 36 386, EP-A 854 129, US-A 4,317,926 , DE-A 198 37 520, DE-A 196 06 877, DE-A 190 50 1325, DE-A 102 47 240, DE-A 197 40 253, EP-A 695 736, EP- A 1 041 062, EP-A 117 146, DE-A 43 08 087, DE-A 43 35 172, DE-A 44 36 243, DE-A 199 24 532, DE-A 103 32 758 and DE- Quot; condensation of the target product "described in U.S. Patent No. 24 533, and further post-treatment methods of the" condensate ". Acrylic acid removal can also be carried out in accordance with EP-A 982 287, EP-A 982 289, DE-A 103 36 386, DE-A 101 15 277, DE-A 196 06 877, DE-A 197 40 252, DE-A 196 27 847 , EP-A 920 408, EP-A 1 068 174, EP-A 1 066 239, EP-A 1 066 240, WO 00/53560, WO 00/53561, DE-A 100 53 086 and EP- As shown in FIG. As in FIG. 7 of WO / 0196271 or as described in DE-A 10 2004 032 129 and its equivalent patents. Suitable removal methods are also those described in documents WO 04/063138, WO 04/35514, DE-A 102 43 625 and DE-A 102 35 847. Further treatment of the obtained crude acrylic acid can be carried out, for example, as described in documents WO 01/77056, WO 03/041832, WO 02/055469, WO 03/078378 and WO 03/041833.

Typically, the product gas B is supplied to the separation zone B at an operating pressure leaving the reaction zone B.

A common feature of the separation process is that it mainly comprises (as already mentioned) a component of the product gas B (i.e., a component that is difficult to condense or volatilize) at a standard pressure (1 bar) lower than the boiling point water and usually below -30 ° C A stream of residual gas is usually left on top of a particular separation column, for example containing a separating column, wherein the product gas B is fed into the lower compartment, for example, usually after its direct and / or indirect cooling has proceeded . However, the residual gas may also include, for example, still components such as water vapor and acrylic acid.

In the lower compartment of the separation column, the less volatile components of the product gas B, including the particular target product, are mainly obtained on the condensation. The condensed aqueous phase is generally removed via lateral extraction and / or through the bottom.

Typically, the (main) residual gas comprises the following contents:

1 to 20% by volume of H ₂ O,

N₂ 0 to 80% by volume,

10 to 90% by volume of propane,

H₂ 0 to 20% by volume,

O₂ 0 to 10% by volume,

CO₂ 1 to 20% by volume, and

CO 0 to 5% by volume.

The content of acrolein and acrylic acid is generally not more than 1 volume% in each case.

The preferred (main) residual gas is

1 to 5 vol% H ₂ O,

10 to 80% by volume of propane,

0 to 5% by volume of N ₂ ,

0 to 20% by volume of H ₂ ,

1 to 10% by volume of O ₂ ,

0 to 5% by volume of CO, and

1 to 15% by volume of CO ₂ .

Typically, the residual gas is discharged in the separation zone II at a pressure of at least 1 bar and at most 3 bar, preferably at most 2.5 bar, generally at most 2 bar, according to the invention. The temperature of the residual gas is typically less than or equal to 100 ° C, typically less than or equal to 50 ° C (but typically greater than or equal to 0 ° C).

Suitably from an application point of view, the propane present in the residual gas is suitably absorbed in the separation zone III at a pressure of from 5 to 50 bar, preferably from 10 to 25 bar. In the compression zone, the residual gas leaving the separation zone II will usually be compressed to a pressure favorable to absorption of the propane (which is advantageously carried out in several stages). Preferably, in accordance with the present invention, the compression of the residual gas is carried out in a first stage of compression, for example in a turbo compressor (radial compressor, for example a MH4B type radial compressor from Mannesmann DEMAG, Germany) (For example, 1.20 bar to 2.40 bar; generally, compression is effected in the first compression stage in the reaction zone A using a gas compressor, also referred to as a gas compressor), essentially at an operating pressure corresponding to the working pressure in the reaction zone A. With a pressure somewhat exceeding the working pressure used). This heats the residual gas to, for example, 50 캜 or lower. If not the total residual gas provided to the propane absorption, the compressed residual gas, as described, is suitably divided into two parts of the same composition. The proportion of the smaller portion can be, for example, 49% or less, preferably 40% or less, or 30% or less, or 20% or less, or 10% or less of the total amount. It may also be at least 10% and at least 49% of the total amount.

This portion is available immediately as a residual gas circulating gas to be recycled to the reaction zone A. The other part is further compressed with the working pressure considered for propane absorption. This can in principle be carried out in only one additional compression step. Advantageously, however, according to the present invention, two or more additional compression stages will be used for this purpose. This is due to the fact that the compression of the gas (as already indicated in the first compression step) is associated with its temperature increase. Conversely, the decompression (expansion) of a "propane-scrubbed" residual gas (where appropriate provided for H ₂ film removal) is associated with its cooling. If this depressurization is likewise carried out in several stages, indirect heat exchange between the heated, compressed, non-scrubbed residual gas and cooling, expansion, and "propane-scrubbed" residual gas can be performed in each case before the next step. This heat exchange (which may additionally be supported by indirect air cooling of the compressed residual gas) can also be used prior to the first expansion (as a result of indirect heat exchange, the condensation of any water vapor still present in the residual gas, It can happen inside. Advantageously, from an application point of view, said expansion is carried out in an expansion turbine (advantageously also in a multistage manner) (this is used for the recovery of the compression energy; in other words, the mechanical energy obtained in the expansion, Ex) can be used directly as an additional or main power to one of the compressors, or it can be used to generate electricity). Depending on the compression or expansion step, the difference between the inlet and outlet pressures may be, for example, 2 to 15 bar.

Before the propane present in the compressed residual gas is absorbed into the organic solvent (scrubbed from the residual gas compressed by the organic solvent (absorbent)) from the residual gas compressed in the separation zone III, May be provided in an additional separation zone for separation to remove at least a portion of any hydrogen molecules still present therein. Suitable separation membranes in this context are, for example, aromatic polyimide membranes, such as UBE Industries Ltd. Of these, membrane types A, BH, C and D are particularly useful. For this removal, it is particularly preferred to use a BH polyimide film from UBE Industries. For this membrane, the hydrogen permeability for H ₂ at 60 ° C is 0.7 · 10 ^-3 [STP · cc / cm ² · sec · cm Hg]. For this purpose, the compressed residual gas is generally tubular (plate-type or winding modules are also useful) and can pass through membranes that are permeable only to hydrogen molecules. As a result, the removed hydrogen molecule may be used in addition to other chemical syntheses or at least partially recycled to reaction zone A (suitably as a component of the reaction gas A charge gas mixture).

Membrane separation of hydrogen molecules from the compressed residual gas is suitable according to the invention because it is preferably also carried out under high pressure (for example 5 to 50 bar, typically 10 to 25 bar). Thus, according to the present invention, the hydrogen membrane separation which is carried out immediately before the propane absorption (propane scrubbing) from the residual gas, or alternatively immediately after the propane scrubbing (propane absorption) of the residual gas (preferably according to the invention) , It is advantageous to use twice the pressure level once increased in an advantageous manner. In addition to the plate membrane, a wound-up membrane or a capillary membrane or a tubular membrane (hollow fiber membrane) is particularly useful for hydrogen removal. Their inner diameter may be, for example, from several micrometers to several millimeters. For this purpose, similar to the tube of the vortex reactor, for example, the tube bundle is cast into one plate at each end of the tube in each case. Outside the tube, preferably a reduced pressure (less than 1 bar) prevails. The residual gas under pressure (or "propane-scrubbed" residue gas) is delivered to one of the two plate ends and is conveyed through the tube interior to the tube outlet at the opposite plate end. As a result, along the flow path defined within the tube, hydrogen molecules are released out through the H ₂ permeable membrane.

There is essentially no restriction on the performance of the propane absorption from the remaining residual gas (the unconverted propylene is generally absorbed with propane during the course of partial oxidation in reaction zone B). A simple process for this purpose is, for example, to remove residual gas (suitably according to the invention, the temperature is 10 to 100 DEG C, preferably 10 to 70 DEG C) at a pressure of 5 to 50 bar, At a pressure of 25 bar, the temperature advantageously ranges from 0 to 100 占 폚, preferably from 20 to 50 占 폚 or 40 占 폚, and propane and propylene are passed (for example simply passing) (Preferably hydrophobic) absorbing organic solvent (preferably hydrophobic)

For example, a gas (e. G., N ₂ ) that behaves inertly to reaction zone A and / or a gas required as a reactant in this reaction zone (e. G., Air or a mixture of another oxygen molecule and an inert gas Desorption (rectification, evaporation), rectification and / or stripping with water, preferably steam, or a mixture of water vapor and oxygen molecules, or a mixture of water vapor and air, And recycled to reaction zone A as a propane-containing feed stream.

Suitable absorbents for absorption and removal are in principle all absorbents capable of absorbing propane and propylene. The sorbent is preferably a hydrophobic and / or high boiling organic solvent. Advantageously, the boiling point of the solvent (at a standard pressure of 1 bar) is at least 120 ° C, preferably at least 180 ° C, predominantly at from 200 to 350 ° C, especially from 250 to 300 ° C, more preferably from 260 to 290 ° C . Suitably, the flashpoint (at standard pressure 1 bar) is above 110 ° C. Generally suitable as the absorbent are organic solvents which are relatively nonpolar, for example aliphatic hydrocarbons which preferably do not contain any external active polar groups, and also aromatic hydrocarbons. In general, it is preferred that the absorbent has a very high solubility simultaneously for propane and propylene and has a very high boiling point. Examples of absorbents include aliphatic hydrocarbons such as C ₈ -C ₂₀ - alkane or alkene, an aromatic hydrocarbon, such as paraffin-ether intermediate having a group five days fraction or bulky (sterically are required) on the oxygen atom from the distillation , Or a mixture thereof, to which a polar solvent, for example dimethyl 1,2-phthalate as disclosed in DE-A 43 08 087, may be added. Also suitable are esters of benzoic acid and phthalic acid and straight chain alkanols containing from 1 to 8 carbon atoms, such as n-butyl benzoate, methyl benzoate, ethyl benzoate, dimethyl phthalate, diethyl phthalate, For example, diphenyl, diphenyl ether, mixtures of diphenyl and diphenyl ethers or chlorine derivatives thereof, and triarylalkenes such as 4-methyl-4'-benzyldiphenylmethane and its isomers , 2-methyl-2'-benzyldiphenylmethane, 2-methyl-4'-benzyldiphenylmethane and 4-methyl-2'-benzyldiphenylmethane, and mixtures of the above isomers. Suitable absorbents are the azeotropic compositions of solvent mixtures of diphenyl and diphenyl ethers, preferably in particular about 25% by weight of diphenyl (biphenyl) and about 75% by weight of diphenyl ether, for example commercially available diphenyls (E.g., Bayer Aktiengesellschaft). Often, such a solvent mixture comprises a solvent such as dimethyl phthalate added in an amount of 0.1 to 25% by weight based on the total solvent mixture. Particularly suitable absorbents are also octane, nonane, decane, undecane, dodecane, tridecane, tetradecane, pentadecane, hexadecane, heptadecane and octadecane, with tetradecane being particularly suitable. It is preferable that the used absorbent first meet the above-mentioned boiling point and secondly does not have a too high molecular weight at the same time. Advantageously, the molecular weight of the sorbent is less than 300 g / mol. Also suitable are paraffin oils having 8 to 16 carbon atoms, as described in DE-A 33 13 573. Examples of suitable commercial products include products commercially available from Haltermann, such as Halpasol i, such as Halfasol 250/340 i and Halfasol 250/275 i, There is a printing ink oil under Printosol. Commercial products that do not contain aromatics, such as products of the PKWFaf type, are preferred. If they contain a small amount of residual aromatics, they may be advantageously reduced prior to use by rectification and / or adsorption and may be lowered significantly to a value of less than 1000 ppm by weight. Additional suitable commercially available products include n- paraffin (C ₁₃ -C ₁₇₎ or Toulon be - Raffine re-Ems land Michaela goal geem beha (Erdoel-Raffinerie-Emsland GmbH) (Mihagol, registered trademark), Conde Ah Augusta LINPAR (registered trademark) of SASOL Italy SpA, and SLOVNAFT of Slovakia), which are common in CONDEA Augusta SpA (Italy) or SASOL Italy SpA there are a paraffin (in) C ₁₄ -C _18.

In the above-mentioned linear hydrocarbon products, the content (recorded by area percentage of gas chromatography analysis) is typically as follows:

Total C ₉ to C ₁₃ : less than 1%; C ₁₄ : 30 to 40%; C ₁₅ : 20 to 33%; C ₁₆ : 18 to 26%; C ₁₇ : 18% or less; C _{≥ 18} : less than 2%.

A typical composition of a zeol product is as follows:

C ₁₃ : 0.48%; C ₁₄ : 39.8%; C ₁₅ : 20.8%; C ₁₆ : 18.9%; C ₁₇ : 17.3%; C ₁₈ : 0.91%; C _19: 0.21%.

Typical compositions of Haltermann products are as follows:

C ₁₃ : 0.58%; C ₁₄ : 33.4%; C _15: 32.8%; C ₁₆ : 25.5%; C ₁₇ : 6.8%; C _{≥ 18} : less than 0.2%.

In continuous operation, the composition of the absorbent will vary correspondingly as a result of the process.

Absorption can be performed in a column or a quenching apparatus. It is possible to work in parallel or (preferably) countercurrent. Suitable has an absorption column, for example - of (bubble cap and / or having a sieve tray) column (e. G. A specific surface area 100 has a tray column, a structured packing to 1000 m ² / m ³ or 750 m ² / m ³ Sheet metal packing, such as Mellapak (R) 250 Y) and columns with random packing (e.g. filled with Raschig packing). However, it is also possible to use trickle and spray towers, graphite block sorbents, surface sorbents such as thick and thin sorbents, and also plate scrubbers, cross-spray scrubbers and rotary scrubbers. It may also be desirable to allow the absorption to take place in the internal water-containing and non-containing bubble columns.

Propane and, if appropriate, propylene can be removed by stripping, flashing and / or distillation (rectification) from the absorbent (absorbent).

Propane and propylene are preferably removed by stripping and / or desorption from the absorbent. The desorption can be effected by pressure and / or temperature changes, for example at a pressure of from 0.1 to 10 bar, or from 1 to 5 bar, or from 1 to 3 bar and at a temperature of from 0 to 200 ° C, especially from 20 to 100 ° C, Deg.] C to 70 [deg.] C, particularly preferably 30 to 50 [deg.] C. A preferred stripping gas for stripping is water vapor. A useful ophthalmic alternative is a mixture of nitrogen molecules or nitrogen molecules and water vapor. However, a mixture of carbon dioxide or water vapor and / or nitrogen is also a suitable stripping gas according to the present invention. Particularly suitable for stripping is also an internal water-containing and free bubble column.

Propane and, if appropriate, propylene can also be removed by distillation or rectification from the absorbent (absorbent), in which case a column familiar with the skilled person and having structured packing, random packing or corresponding internals can be used. Possible conditions for the distillation or rectification are a pressure of 0.01 to 5 bar, or 0.1 to 4 bar, or a pressure of 1 to 3 bar and a (lower) temperature of 50 to 300 ° C, especially 150 to 250 ° C.

Before being used to fill the reaction zone A, the propane circulating gas obtained by stripping from the absorbent may be sent to further processing steps in the separation zone IV, for example to reduce the loss of entrained sorbent (e. G. (demister) and / or separation in a deep-seated filter), so that the reaction zone can be protected from the absorbent at the same time. Removal of such an absorbent can be accomplished by any process modification known to those skilled in the art. An example of such a preferred removal embodiment in the process of the present invention is to quench the stream discharged from the stripping apparatus with water. In this case, the entrained effluent stream in the absorbent is washed away with water, and the effluent stream is simultaneously loaded with water. This washing, or quenching, can be performed, for example, by a liquid collection tray at the top of the desorption column, by back spraying of water, or in a dedicated device. In order to facilitate the separation effect, it is possible to provide an internal water which increases the quenching surface area in the quenching chamber, as is known to those skilled in the art from rectification, adsorption and desorption.

Decompression of the absorbent comprising propane (and, if appropriate, propylene) to a pressure suitable for the removal of propane (and propylene if appropriate) can be carried out, for example, by a valve or a reverse pump in the process according to the invention. In the case of an inverse pump, the released mechanical energy is suitably used according to the invention to recompress the absorber from which propane (and, if appropriate, propylene) has been removed (for example in a desorption column or stripping column). In general, the propane-free absorbent can be reused as an absorbent (suitably recompressed by the pump at the adsorption pressure according to the invention) to remove the propane from the residual gas in the separation zone III.

Particularly advantageously according to the invention, propane is separated from the absorption in the separation zone IV by 1.5 or 2 to 5 bar, preferably from 2 to 5 bar, more preferably from 2.5 to 4.5 bar, most preferably from 3 to 4 bar bar pressure.

The removal of propane from the absorbent at the above-mentioned working pressure is very particularly advantageous because the propane circulating gas obtained at the working pressure can be directly recycled to the reaction zone A. This propane circulating gas can be used as a driving jet for the dehydrogenation circulating gas flow by means of the jet pump principle according to the teaching of DE-A 102 11 275. Suitably according to the invention, the propane circulating gas is preheated to a suitable temperature in the reaction zone A by indirect heat exchange with the hot reaction gas A and / or the hot product gas A in advance.

The typical content of propane circulating gas obtained by desorption and / or steam stripping from the propane-containing absorbent in the separation zone IV may be as follows in the process according to the invention:

80 to 99.99 mol% propane,

0 to 5 mol% of propylene, and

H ₂ O 0 to 20 mol%.

In many cases, the content of propane circulating gas is as follows:

80 to 99 mol% of propane,

1 to 4 mol% of propylene, and

H ₂ O 2 to 15 mol%.

In addition to the new propane (as a component of crude propane) and propane circulating gas and, if appropriate, dehydrogenated recycle gas, a residual gas recycle gas (suitably compressed) and / or one (other) A, the crude propane and the residual gas circulating gas and / or the oxygen molecule-containing gas are suitably mixed in advance from an application point of view to provide a first gas mixture, and this first gas mixture (After being heated by indirect heat exchange with the high temperature reaction gas A and / or the product gas A) is preferably a mixture of a propane circulating gas or a dehydrogenated recirculating gas (preferably an indirect heat of the high temperature reaction gas A and / ), And the product gas mixture is passed to the reaction zone A as the reaction gas A charge gas mixture The. Advantageously, hydrogen molecules are also added to the resulting gas mixture, which is then fed into reaction zone A as reaction gas A charge gas mixture.

The requirements for the oxygen molecule-containing gas in the process according to the present invention are suitably taken from (and covered from) conventional sources from an application point of view.

In accordance with the present invention, crude propane is evaporated by condensation (e. G. Acidic water) obtained in separation zone II, suitably by indirect heat exchange (so that the cooled condensate is condensed to obtain a new condensate Can be used for direct cooling in separation zone II). Alternatively, other liquid phases (suitably at a temperature of 20 to 40 占 폚, preferably 25 to 35 占 폚) may also be used as thermal carriers for indirect heat exchange for the purpose of vaporizing crude propane. In order to further heat the crude propane, indirect heat exchange with the hot water vapor, which is typically obtained as a by-product in the process according to the invention (for example during the removal of the process heat obtained during the partial oxidation process) bar and 152 < 0 > C) are generally used.

If necessary, a small amount of the absorbent is continuously released in the separation zone IV and replaced by a new absorbent. The released absorbent may be post-treated by rectification to provide a fresh absorbent. The residual gas, which is compressed as described and pre-"propane-scrubbed" and then, if appropriate, membrane hydrogenated, is typically sent to residue conversions (offgas conversions). Furthermore, the high boiling point water removed in the separation zone II and the aqueous condensate (for example acidic water) formed between the separation zone II and, where appropriate, the separation zone II and the separation zone III, are supported, if appropriate, Which is carried out using air as a source for < RTI ID = 0.0 > a < / RTI > The hot offgas of this painting is typically used for indirect heat exchange with water to produce water vapor, which is then generally discharged to the atmosphere. Typically, it consists exclusively of oxygen molecules, nitrogen molecules, hydrogen molecules and carbon dioxide. Finally, it should be emphasized that the same polymerization inhibitor is added whenever a condensation phase comprising acrylic acid and / or acrolein occurs in the process according to the invention. Useful polymerization inhibitors useful in principle are all known process inhibitors. For example, methylethers of phenothiazine and hydroquinone are particularly suitable according to the present invention. The presence of oxygen molecules increases the effectiveness of polymerization inhibitors.

The acrolein obtained by the process according to the invention can be converted into the acrolein follow-up product mentioned in documents US-A 6,166,263 and US-A 6,187,963. These include 1,3-propanediol, methionine, glutaraldehyde and 3-picoline.

The method according to the invention having the maximum working pressure in the separation zone III is preferred.

Advantageously, for a working pressure P (in each case measured at the inlet to a particular zone) in different zones of the method according to the invention, the following correlation (relational expression) applies:

P _{reaction zone A} > P _{separation zone I} > P _{reaction zone B} > P _{separation zone II} <P _{separation zone III} > P _{separation zone IV} > P _{reaction zone A.}

The separation zone I may also be absent.

In addition, the method according to the present invention can maintain the temperature of the reaction gas A, especially when the reaction gas A contains oxygen molecules at 460 DEG C or lower, preferably 440 DEG C or lower, outside the catalyst and / or the inert layer, Which is advantageous for minimizing the undesired combustion reaction and pyrolysis.

Examples and Comparative Examples

I. Inhomogeneous catalysis of propylene to acrylic acid in the absence and presence of hydrogen molecules Long-term operation of two-stage partial oxidation

A) General experimental setup of reaction equipment

The reactor for the first oxidation step

The reactor consisted of a stainless steel jacketed cylinder (the cylindrical guide tube was enclosed by a cylindrical outer container). The wall thickness was always 2 to 5 mm.

The inner diameter of the outer cylinder was 91 mm. The inner diameter of the guide tube was approximately 60 mm.

At the top and bottom, the jacket-shaped cylinder was terminated with a lid and a base respectively.

The catalytic tube (total length 400 cm, inner diameter 26 mm, outer diameter 30 mm, wall thickness 2 mm, stainless steel) was extruded through the lid and base at its higher or lower end And received in a guide tube. A heat exchange medium (salt melt consisting of 53 wt% potassium nitrate, 40 wt% sodium nitrite and 7 wt% sodium nitrate) was enclosed in a cylindrical vessel. The heat exchange medium was cyclically pumped by a propeller pump to ensure a very homogeneous thermal boundary condition on the outer wall of the catalyst tube over the entire length of the catalyst tube in a cylindrical vessel (400 cm).

An electric heater attached to the outer jacket adjusted the temperature of the heat exchange medium to the desired level. Otherwise, there was air cooling.

Reactor charge:

First stage reactor As seen above, the salt melt and the charge gas mixture of the first stage reactor were delivered in stream. The charge gas mixture was introduced into the first-stage reactor from the bottom. This was transferred to a reaction tube having a temperature of 165 ° C in each case.

The salt melt was introduced into the cylindrical guide tube from the bottom of the temperature T ⁱⁿ and flowed ^{out to} the top of the temperature T ^{out in} the cylindrical guide tube where T ^out was 2 ° C below T ⁱⁿ .

The T ⁱⁿ was adjusted so that at the outlet of the first oxidation step, a propylene conversion of 97.8 ± 0.1 mol% always occurred in a single pass.

Catalyst tube charging: (bottom to top)

Compartment A: Length 90 cm

Upper layer of a bedrock having a diameter of 4 to 5 mm.

Compartment B: Length 100 cm

Catalytic charging of a homogeneous mixture of 30% by weight of a continuous rock having dimensions of 5 mm x 3 mm x 2 mm (outer diameter x length x inner diameter) and 70% by weight unsupported catalyst from compartment C.

Compartment C: Length 200 cm

(5 mm x 3 mm x 2 mm = outer diameter x length x inner diameter) unsupported catalyst according to Example 1 of DE-A 100 46 957 (stoichiometry: [Bi ₂ W ₂ O ₉ x 2WO ₃ ] _0.5 [ Mo ₁₂ Co _5.5 Fe _2.94 Si _1.59 K _0.08 O _x ] ₁ ).

Compartment D: Length 10 cm

Dimension Downstream layer of a sheet of 7 mm × 3 mm × 4 mm (outer diameter × length × inner diameter).

Intermediate cooling and intermediate oxygen supply (pure O ₂ as secondary gas)

For intermediate cooling (indirectly using air), the product gas mixture exiting the first fixed bed reactor was loaded centrally for a length of 20 cm and had dimensions of 7 mm x 3 mm x 4 mm (outer diameter x length x (Length: 40 cm, inner diameter: 26 mm, outer diameter: 30 mm, wall thickness: 2 cm, stainless steel, insulated material: 1 cm) Surrounding it).

The product gas mixture was always introduced into the connection tube at a temperature > T ⁱⁿ (step 1) and allowed to flow out at a temperature above 200 ° C and below 270 ° C.

At the end of the connector, oxygen molecules were metered into the cooled product gas mixture at the pressure level of the product gas mixture. The resulting gas mixture (the charge gas mixture for the second oxidation step) was passed directly to the flanged second-stage catalyst tube by the other end of the connector. The amount of metered molecular oxygen was such that the molar ratio of acrolein present in the resulting gas mixture to the O ₂ present in the resulting gas mixture was 1.3.

The reactor for the second oxidation step

A catalyst tube fixed bed reactor, the same design as the first oxidation step, was used. The salt melt and charge gas mixture was delivered in parallel as viewed from the reactor. The salt melt was introduced from the bottom, and the charge gas mixture was also carried out. The inlet temperature T ⁱⁿ of the salt melt was adjusted so that an acrolein conversion of 99.3 + 0.1 mol% was always produced in a single pass at the outlet of the second oxidation stage. The T ^out of the salt melt was 2 ° C higher than T ⁱⁿ .

Catalyst tube charging (from bottom to top):

Compartment A: Length 70 cm

Dimensions A preliminary layer of a lamella of 7 mm × 3 mm × 4 mm (outer diameter × length × inner diameter).

Compartment B: Length 100 cm

Catalytic charging of a homogeneous mixture of 30% by weight of a talcite ring of dimensions 7 mm x 3 mm x 4 mm (outer diameter x length x inner diameter) and 70% by weight of the coating catalyst from zone C.

Compartment C: Length 200 cm

(7 mm x 3 mm x 4 mm = outer diameter x length x inner diameter) coated catalyst according to Production Example 5 of DE-A 10046928 (stoichiometry: Mo ₁₂ V ₃ W _1.2 Cu _2.4 O _x ).

Compartment D: Length 30 cm

Downstream layer of a bedrock having a diameter of 4-5 mm.

B) The result is achieved as a function of the composition of the packed gas mixture in the first oxidation step (the propene loading is set at 150 l (STP) / l · h; the selectivity of acrylic acid formation is always at least 94 mol% negative).

a) the composition of the charge gas mixture for the first oxidation step was substantially as follows:

7% by volume of propylene,

12 vol% O ₂ ,

20% by volume of H ₂ ,

5% by volume of H ₂ O, and

56 vol% N _2.

The inlet temperature at the start of the fresh charged reactor was as follows:

T ⁱⁿ (first oxidation step): 320 ° C

T ⁱⁿ (second oxidation step): 268 ° C

After 3 months of operation time, the influent temperature was as follows:

T ⁱⁿ (first oxidation step): 330 ° C

T ⁱⁿ (second oxidation step): 285 ° C

b) the composition of the packed gas mixture for the first oxidation step was substantially as follows:

7% by volume of propylene,

12 vol% O ₂ ,

5% by volume of H ₂ O, and

76 vol% N ₂ .

The inlet temperature at the start of the fresh charged reactor was as follows:

T ⁱⁿ (first oxidation step): 320 ° C

T ⁱⁿ (second oxidation step): 268 ° C

After 3 months of operation time, the influent temperature was as follows:

T ⁱⁿ (first oxidation step): 324 ° C

T ⁱⁿ (second oxidation step): 276 ° C

II. The first exemplary method for producing acrylic acid from propane (described as normal operation)

The reaction zone A consisted of a reactor with a shaft designed as a tray reactor, adiabatically constructed, and had three fixed catalyst beds arranged continuously in the flow direction. The particular fixed catalyst bed consists of a layer of inert material (layer height: 26 mm; diameter 1.5 to 2.5 mm) arranged on the stainless steel wire mesh (arranged in the order specified in the direction of flow) and a new dehydrogenation catalyst, (Diameter of 1.5 to 2.5 mm) (dehydrogenation catalyst: bed volume ratio of bedrock of 1: 3). (Alternatively, the same amount of catalyst may be used undiluted at this point).

A fixed gas mixer was placed upstream of each fixed bed. The dehydrogenation catalyst is catalyzed by elemental oxidized forms of Cs, K and La and is extruded from a ZrO ₂ .SiO ₂ mixed oxide support extrudate (average length (Gaussian distribution in the range of 3 to 12 mm at maximum about 6 mm): 6 mm, Pt alloy coated with the element stoichiometry (mass ratio including the support) of Pt _0.3 Sn _0.6 La _3.0 Cs _0.5 K _0.2 (ZrO ₂ ) _88.3 (SiO ₂ ) _7.1 on the outer and inner surfaces of the Pt / Sn alloy (Preparation of catalyst precursor as in Example 4 of DE-A 10219879 and activation of the active catalyst).

Inhomogeneously catalyzed partial propane dehydrogenation was carried out using a straight-through, dehydrogenation recycle gas mode in the tray reactor. The propane loading for all trays in all trays (calculated without inert materials) was 1500 l (STP) / l · h.

A reaction gas A charge gas mixture (starting mixture; P = 2.13 bar) of 7.30 m ³ (STP) / h was fed to the first catalyst bed in the flow direction and the content thereof was as follows:

61.22% by volume of propane,

12.17% by volume of propylene,

0.86% by volume of ethylene,

2.56% by volume of methane,

3.78 vol.% H ₂ ,

O ₂ 2.10% by volume,

12.08% by volume of H ₂ O,

0.38% by volume of CO, and

CO ₂ 3.35% by volume.

Reaction gas A The charge gas mixture is a mixture of a first gas mixture (2.34 m ³ (STP) / h) and a second gas mixture (4.96 m ³ (STP) / h) &Lt; / RTI &gt; The temperature of the first gas mixture was 400 DEG C and the pressure was 2.13 bar, consisting of:

1.19 m ³ (STP) / h of compressed residual gas circulating gas (T = 70 ° C, P = 2.40 bar)

1.05 m ³ (STP) / h crude propane (T = 237 ° C, P = 2.90 bar) containing propane (fresh propane) to an extent greater than 98%

0.098 m ³ (STP) / h of O ₂ (T = 139 ° C, P = 2.33 bar, purity greater than 99% by volume).

In the first heat exchanger, the indirect heat exchange with the product gas A delivered in the direction of the reaction zone B outside the reaction zone A set the temperature of 400 ° C. The second gas mixture was constructed as follows.

Dehydrogenation circulating gas of 3.68 m ³ (STP) / h (T = 551 ° C, P = 2.05 bar), and

Propane circulating gas of 1.28 m ³ (STP) / h (T = 400 ° C, P = 3.26 bar).

Its pressure was 2.13 bar and the temperature was 506 ° C.

Indirect heat exchange with the product gas A cooled to a temperature of 435 캜 in the first heat exchanger set the temperature of the propane circulating gas. Dehydrogenation The circulating gas flow was carried out in accordance with the jet pump principle (DE-A 10211275) and the propane circulating gas functioning as the drive jet was heated to 400 ° C.

The inlet temperature of the Reaction Gas A packed gas mixture and the bed height of the first catalyst layer through which the Reaction Gas A packed gas mixture flows were such that the Reaction Gas A was spilled out in the catalyst bed at a temperature of 515 ° C and a pressure of 2.11 bar in the following amounts:

55.8% by volume of propane,

15.85% by volume of propylene,

0.88% by volume of ethylene,

2.89% by volume of methane,

3.78 vol.% H ² ,

O ₂ 0% by volume,

15.5% by volume of H ₂ O, and

CO ₂ 3.67% by volume.

The effluent was 7.46 m ³ (STP) / h.

Over the first catalyst bed, 0.12 m ³ (STP) / h of oxygen molecules (purity greater than 99% by volume) were metered into the reaction gas A (T = 139 ° C, P = 2.33 bar). The metering addition was carried out by throttling so that the produced pressure of the reaction gas A produced was still 2.11 bar.

The layer height of the second catalyst layer was such that the reactive gas A was spilled out in the second catalyst bed at a temperature of 533 DEG C and a pressure of 2.08 bar in the following amounts:

49.2% by volume of propane,

19.24% by volume of propylene,

1.04% by volume of ethylene,

3.14% by volume of methane,

H ₂ 4.43% by volume,

O ₂ 0% by volume,

17.98% by volume of H ₂ O, and

3.52% by volume of CO ₂ .

The effluent was 7.79 m ³ (STP) / h.

At the upstream of the stationary mixer located upstream of the third catalyst bed, 0.13 m ³ (STP) / h of oxygen molecules (purity greater than 99% by volume) were metered into this reaction gas A (T = 139 ° C, pressure 2.33 bar Lt; / RTI &gt; The pressure of the produced reaction gas A was still 2.08 bar.

The layer height of the third catalyst bed was such that the reaction gas A was spilled out as product gas A in the third catalyst bed at a temperature of 551 DEG C and a pressure of 2.05 bar in the following amounts:

42.19% by volume of propane,

22.79% by volume of propylene,

1.14% by volume of ethylene,

3.42% by volume of methane,

H ₂ 5.28% by volume,

O ₂ 0% by volume,

20.32% by volume of H ₂ O, and

CO ₂ 3.36% by volume.

The effluent was 8.17 m ³ (STP) / h.

The product gas A was divided into two parts of the same composition by a flow divider. The first part was 3.68 m ³ (STP) / h and the second part was 4.49 m ³ (STP) / h. The first portion was recycled to the reaction zone A as a component of the reaction gas A charge gas mixture. For the dehydrogenation circulating gas flow of the first part, the propane circulating gas heated as above is used as the driving jet to operate the jet pump, and the delivery direction of the driving jet compressed through the driving nozzle through the mixing zone and the diffuser is The suction nozzle refers to the direction of the first flow portion of the product gas A and the "suction nozzle-mixing zone-diffuser" link connects the first portion of the product gas A to be recycled and the first portion of the reaction gas A &lt; / RTI &gt; A second portion of the product gas A is delivered out of the reaction zone A and a first portion of the first gaseous mixture comprised of the reactive gas A starting mixture (charge gas mixture) and consisting of the compressed residual gas circulating gas, crude propane, Lt; RTI ID = 0.0 &gt; 555 C &lt; / RTI &gt; by indirect heat exchange. By doing this, the working pressure in the product gas A delivered out of the reaction zone A dropped from 2.05 bar to 1.95 bar. At the same time, the first gaseous mixture was heated to 16O &lt; 0 &gt; C at 16O &lt; 0 &gt; C and pressure dropped from 2.40 bar to 2.13 bar.

In the second indirect indirect heat exchange with the propane circulating gas from the separation zone IV, the product gas A was cooled from 435 캜 to 335 캜. By doing this, the working pressure in product gas A dropped to 1.94 bar. Conversely, the temperature of the propane circulating gas increased from 69 ° C to 400 ° C.

Thereafter, water vapor of 0.76 m ³ (STP) / h was removed by condensation from product gas A cooled to 335 ° C. For this, the product gas A cooled to 335 ° C is first subjected to indirect heat exchange to produce product gas A ^* (T = 40 ° C, P = 1.68 bar) of 3.73 m ³ (STP) / h Cooled to &lt; RTI ID = 0.0 &gt; 108 C). &Lt; / RTI &gt; The temperature of the product gas A ^* was simultaneously increased to 305 ° C. In the downstream air cooler, the product gas A (T = 108 캜, P = 1.94 bar) was cooled to 54 캜 by indirect heat exchange with outside air (the working pressure of product gas A dropped to 1.73 bar) Thereafter, it had two phases.

(29 [deg.] C) and directly removed from the product gas A by appropriate direct cooling, removes the aforementioned amount of water vapor from the product gas A by condensation to produce product gas A ^* (P = 1.68 bar, T = 40 캜, 3.73 m ³ (STP) / h). This was heated to 305 DEG C as described above. Then, oxygen molecules of 1.39 m ³ (STP) / h (purity greater than 99% by volume, T = 139 ° C, P = 2.33 bar) were throttled with product gas A ^* . This formed a reaction gas B starting mixture (charge gas mixture) having a temperature of 288 占 폚 and a working pressure of 1.63 bar.

Reaction gas B The starting mixture (charge gas mixture) had the following contents:

37.01% by volume of propane,

19.99% by volume of propylene,

0.99% by volume of ethylene,

2.99% by volume of methane,

H ₂ 4.63% by volume,

O ₂ 26.79% by volume,

3.14% by volume of H ₂ O, and

CO ₂ 2.95% by volume.

Reaction zone B was charged using this.

The first oxidation step in the flow direction of the reaction gas B was a multi-tube reactor having two temperature zones. The reaction tube was constructed as follows: V2A steel; Outer diameter 30 mm, wall thickness 2 mm, inner diameter 26 mm, length 350 cm. From top to bottom, the reaction tube was charged as follows:

Compartment 1: Length 50 cm

As the upper layer, there is a circle of dimensions of 7 mm × 7 mm × 4 mm (outer diameter × length × inner diameter).

Compartment 2: Length 140 cm

A homogeneous homogeneous mixture of 20% by weight (alternatively 30% by weight) of a 5 mm × 3 mm × 2 mm (outer diameter × length × inner diameter) Catalytic charging of the mixture.

Compartment 3: Length 160 cm

(5 mm x 3 mm x 2 mm = outer diameter x length x inner diameter) unsupported catalyst according to Example 1 of DE-A 100 46 957 (stoichiometry: [Bi ₂ W ₂ O ₉ .2WO ₃ ] _0.5 [Mo ₁₂ Co _5.5 Fe _2.94 Si _1.59 K _0.08 O _x ] ₁ ).

Alternatively, it is also possible to use one of the catalysts EUC 1 to EUC 11 from investigation starting document No. 497012 (Aug. 29, 2005) (the specific surface area of the active composition is reported in cm ² / g The exact dimensions are m ² / g with the same numerical value).

From the top to the bottom, the first 175 cm was temperature controlled by salt bath A pumped countercurrent to reaction gas B. The second 175 cm was temperature controlled by salt bath B pumped countercurrent to reaction gas B.

The second oxidation step in the flow direction of the reaction gas B was also a multi-tube reactor with two temperature zones. The reaction tube was charged from top to bottom as follows:

Compartment 1: Length 20 cm

As the upper layer, there is a circle of dimensions of 7 mm × 7 mm × 4 mm (outer diameter × length × inner diameter).

Compartment 2: Length 90 cm

A homogeneous mixture of 25% by weight (alternatively 30% by weight) of a 7 mm × 3 mm × 4 mm (outer diameter × length × inner diameter) copper oxide ring and 75% by weight of coating catalyst from Zone 4 (alternatively 70% by weight) Lt; / RTI &gt;

Compartment 3: Length 50 cm

A homogeneous mixture of 15 wt.% (Alternatively 20 wt.%) Of the dimensions of 7 mm x 3 mm x 4 mm (outer diameter x length x inner diameter) and 85 wt.% (Alternatively 80 wt.%) Of the coating catalyst from compartment 4 Lt; / RTI &gt;

Compartment 4: Length 190 cm

Catalytic charging of annular (7 mm x 3 mm x 4 mm = outer diameter x length x inner diameter) coated catalyst (stoichiometry: Mo ₁₂ V ₃ W _1.2 Cu _2.4 O _x ) according to Example 5 of DE-A 100 46 928.

From the top to the bottom, the first 175 cm was temperature controlled by Salt C pumped countercurrent to the reaction gas. The second 175 cm temperature was controlled by a salt bath D pumped back into the reaction gas.

In both oxidation steps, each of the salts was transferred to the reaction tube using a deflector plate in a mandrel manner.

Between the two oxidation stages, a vortex heat exchanger cooled by a salt bath capable of cooling the product gas of the first oxidation stage was placed. A valve was provided upstream of the inlet to the second oxidation stage for the supply of oxygen molecules (greater than 99 volume% purity).

During the first oxidation step the catalyst charge was chosen to be 130 l (STP) / l · h. The salt melt (53 wt.% KNO ₃ , 40 wt.% NaNO ₂ , 7 wt.% NaNO ₃ ) had the following inlet temperature:

T _A = 324 ° C T _B = 328 ° C

T _C = 265 ° C T _D = 269 ° C

Sufficient molecular oxygen (139 ℃, 2.33 bar), the first weighed addition (throttling) the product gas mixture of the oxidation step, the 2 O ₂ in the resulting charge gas mixture for the oxidation step: acrolein was a molar ratio of 0.73. The amount of acrolein loaded during the second oxidation step of the catalyst was 110 l (STP) / l · h. The inlet pressure to the second oxidation step was 1.57 bar. Reaction gas B was effluent at an intermediate cooler at a temperature of 260 占 폚 and the inflow temperature of the packed gas mixture to the second oxidation step was 256 占 폚. The product gas mixture of the first oxidation step had the following contents:

16.84% by volume of acrolein,

1.13% by volume of acrylic acid,

36.88% by volume of propane,

0.99 vol% propylene,

2.99% by volume of methane,

H ₂ 4.61% by volume,

0.99% by volume of ethylene,

O ₂ 4.32% by volume,

23.96% by volume of H ₂ O,

0.847% by volume of CO, and

CO ₂ 4.93% by volume.

Before entering the subsequent cooler, the temperature of the product gas in the first oxidation step was 335 ° C.

The product gas mixture (product gas B) in the second oxidation step had a temperature of 270 DEG C, a pressure of 1.55 bar and also had the following contents:

0.089% by volume of acrolein,

17.02% by volume of acrylic acid,

0.46% by volume of acetic acid,

36.79% by volume of propane,

0.99 vol% propylene,

2.98% by volume of methane,

0.99% by volume of ethylene,

4.59% by volume of H ₂ ,

O2 2.69% by volume,

24.57% by volume of H ₂ O,

1.35% by volume of CO, and

CO ₂ 5.91 vol%

As described in WO 2004/035514, product gas B (5.15 m ³ (STP) / h) was fractionally condensed in a tray column (separation zone II).

As the first fuel, 13.7 g / h of high boiling point water (polyacrylic acid (Michael adduct), etc.) was supplied to the residue painting.

2.80 kg / h of condensed crude acrylic acid having a temperature of 15 캜 and 96.99% by weight of acrylic acid was discharged from the second collecting tray to the tray column on the supply of the product gas B. The suspension crystals were separated from the mother liquor in a hydraulic scrubbing column and the mother liquor was recycled to the condensation column as described in WO 2004/035514, as described in WO 2004/035514, after a small amount of water was added and the suspension crystallized. The purity of the washed suspension crystals is greater than 99.87% by weight of acrylic acid and is suitable for the preparation of water- "superabsorbent" polymers for immediate hygiene applications.

The amount of acidic water condensate discharged into the condensation column from the third collection tray on the feed of product gas B but not recycled to the condensation column was 1.10 kg / h, the temperature was 33 ° C, and had the following contents:

0.80% by weight of acrolein,

4.92% by weight of acrylic acid,

5.73% by weight acetic acid, and

87.0% by weight of water.

This was likewise fed to the residue painting (5.49 m ³ (STP) / h of air supplied to the residue painting as the oxygen source).

At the top of the condensation column, 2.98 m ³ (STP) / h of residual gas was sparged in the separation zone II at a temperature of 33 ° C and a pressure of 1.20 bar in the following amounts:

0.03% by volume of acrolein,

0.02% by volume of acrylic acid,

0.01% by volume of acetic acid,

63.7% by volume of propane,

1.72% by volume of propylene,

6.90 vol.% H ₂ ,

5.16% by volume of methane,

1.72% by volume of ethylene,

O ₂ 4.68% by volume,

2.07% by volume of H ₂ O,

2.34% by volume of CO, and

CO ₂ 10.18% by volume.

In the first compressor stage of the multi-stage radial compressor, the residual gas was compressed from 1.20 bar to 2.40 bar, during which the temperature of the residual gas increased to 70 ° C. The compressed residual gas was then divided into two portions of the same composition. The first part (1.19 m ³ (STP) / h) was fed to the formation of the packed gas mixture for reaction zone A as the (residual) gas recycle gas. The other part (1.79 m ³ (STP) / h) was compressed from 2.40 bar to 6.93 bar in the second compressor stage. This was heated to 130 &lt; 0 &gt; C.

In the indirect heat exchanger, it was cooled to 113 ° C without forming condensate (the coolant was "propane-scrubbed" residual gas (hereinafter also referred to as "off-gas"), Starting at 20 bar in the first expansion stage of the multistage expansion turbine and then cooled and cooled at the same time). This heated the off-gas (0.58 m ³ (STP) / h). In the second expansion step, the off-gas was compressed to 1.10 bar, at which time the temperature was 25 [deg.] C, and then fed to the retentate.

In an indirect air cooler, the compressed residual gas (P = 6.93 bar and T = 113 ° C) was cooled to 59 ° C. In the third compressor stage, it was compressed from 6.93 bar to 20 bar and simultaneously heated to 116 ° C.

In a downstream indirect heat exchanger, it was cooled to 107 ° C without forming a condensate (the coolant was a "propane-scrubbed" residual gas, simultaneously heated to 30 ° C to 86 ° C). In the subsequent additional indirect heat exchanger, the residual gas at 107 DEG C was cooled to 35 DEG C (the coolant is surface water at approximately 25 DEG C; it is always possible to use other cooling water herein instead of surface water); This condensed 29 g / h of water out of the residual gas. This aqueous condensate (preferably separated by a liquid separator) was likewise sent to the residue painting and was then drawn with other mentioned residues (supplying 5.49 m ³ (STP) / h of air, as described above) . The high-temperature painting gas was cooled to produce water vapor (52 bar, 267 ° C, 4.6 kg / h) and released to the surroundings.

Residual gas (20 bar, 35 캜) still containing 1.75 m ³ (STP) / h of propane was fed to the lower compartment of the column with random packing (separation zone III). At the top of the scrubbing column was introduced 12.96 kg / h of industrial tetradecane with an inlet temperature of 30 ° C (pressure 20 bar), PKWF 4/7 af type, Germany, Gas chromatographic analysis provides the following GC area percent composition at start-up (new case):

7.6% of n-tridecane,

47.3% of n-tetradecane,

7.0% of n-pentadecane,

3.2% n-hexadecane,

14.1% of n-heptadecane, and

Total residue 20.7%.

At the top of the scrubbing column, a "propane-scrubbed" residual gas (off-gas) was sparged at a pressure of 20 bar and a temperature of 30 ° C and an amount of 0.58 m ³ (STP)

1.97% by volume of propane,

0.049% by volume of propylene,

15.95% by volume of methane,

5.32% by volume of ethylene,

H ₂ 21.29% by volume,

O ₂ 14.46% by volume,

0.82% by volume of H ₂ O,

7.23% by volume of CO, and

CO ₂ 31.43% by volume.

This off-gas was compressed as described, indirect heat exchanged, and then either fed to the residue or drawn by flash.

15.27 kg / h of absorbent with a pressure of 20 bar and a temperature of 59 ° C were discharged from the bottom of the externally heated or uncooled "propane scrubbing column". It contained 14.79 wt% propane and 0.38 wt% propylene.

The adsorbate was heated to 69 [deg.] C in an indirect heat exchanger prior to delivery to the top of the downstream desorption column. The heat carrier used was a liquid effluent from the desorption column (12.96 kg / h), which at a pressure of 3.26 bar was at a temperature of 80 ° C and still contained 0.09% by weight dissolved propane. The mixture was further compressed to 20 bar by a pump, indirectly cooled to 30 ° C using surface water (25 ° C), and recycled to propane adsorption at the top of the "propane scrubbing column".

After heating to 69 [deg.] C, the absorbent is compressed to a pressure of 3.26 bar (for example by an inverted pump or valve) (in the case of an inverted pump the mechanical energy released is suitably also removed from the desorption column, The liquid effluent of the column), and the resulting two-phase mixture was transferred to the top of the desorption column.

A water vapor of 0.1 kg / h was fed to the desorption column (likewise a column with random packing) counter-current to the absorption from the top of the desorption column from the bottom as a stripping gas at a pressure of 5 bar and a temperature of 15 캜.

A heating coil having a temperature of 152 DEG C, a pressure of 5.00 bar, and an amount of 0.5 kg / h was similarly fed back to the water vapor rising the desorption column was placed in the desorption column.

At the top of the desorption column, a 1.28 m ³ (STP) / h of propane circulating gas at a temperature of 69 ° C and a working pressure of 3.26 bar was drained. After heating to 400 DEG C by indirect heat exchange with a product gas A at a temperature of 435 DEG C, propane circulating gas is used as a driving jet to operate the jet pump to convert the dehydrogenated recycle gas and propane circulating gas into the reaction mixture A The reaction zone A was fed to fill.

Propane circulating gas had the following contents:

88.45% by volume of propane,

2.39% by volume of propylene, and

H ₂ O 8.66% by volume.

III. A second exemplary method of producing acrylic acid from propane (described as normal operation)

The reaction zone A is a tray reactor as described under II. Inhomogeneously catalyzed partial propane dehydrogenation was carried out in the tray reactor using a linear pass-through dehydrogenation recycle gas mode. The propane loading for all trays in all trays (calculated without inert materials) was 1500 l (STP) / l · h.

A reaction gas A charge gas mixture (starting mixture; P = 1.84 bar) of 8.23 m ³ (STP) / h was fed to the first catalyst bed in the flow direction, which had the following contents:

64.20% by volume of propane,

16.0% by volume of propylene,

1.39% by volume of methane,

0.46% by volume of ethylene,

H ₂ 2.76% by volume,

1.06% by volume of O ₂ , and

H ₂ O 13.13% by volume.

Reaction gas A The charge gas mixture is a mixture of a first gas mixture (1.15 m ³ (STP) / h) and a second gas mixture (7.08 m ³ (STP) / h) &Lt; / RTI &gt; The temperature of the first gas mixture was 450 DEG C and the pressure was 1.84 bar, consisting of:

1.06 m ³ (STP) / h crude propane (T = 237 ° C, P = 2.90 bar) containing propane (fresh propane) to an extent of greater than 98%

0.087 m ³ (STP) / h of O ₂ (T = 143 ° C, P = 1.84 bar, purity greater than 99% by volume).

In the first heat exchanger, the indirect heat exchange with the product gas A delivered in the direction of the reaction zone B from outside the reaction zone A set the temperature of 450 ° C.

The second gas mixture was constructed as follows.

5.04 m ³ (STP) / h of dehydrogenation circulating gas (T = 540 ° C, P = 1.73 bar), and

Propane circulating gas of 2.04 m ³ (STP) / h (T = 400 ° C, P = 3.26 bar).

Its pressure was 1.84 bar and the temperature was 486 ° C. Indirect heat exchange with the product gas A cooled to a temperature of 476 ° C in the first heat exchanger set the temperature of the propane circulating gas. Dehydrogenation The circulating gas flow was carried out in accordance with the jet pump principle (DE-A 10211275) and the propane circulating gas functioning as the drive jet was heated to 400 ° C.

The inlet temperature of the Reaction Gas A packed gas mixture and the bed height of the first catalyst layer through which the Reaction Gas A packed gas mixture flows are such that the Reaction Gas A is spilled out from the catalyst bed at a temperature of 503 ° C and a pressure of 1.81 bar in the following amounts:

60.7% by volume of propane,

18.19% by volume of propylene,

1.59% by volume of methane,

0.53% by volume of ethylene,

H ₂ 2.94% by volume,

O ₂ 0 vol%, and

H ₂ O 15.0% by volume.

The effluent was 8.35 m ³ (STP) / h.

Over the first catalyst bed, 0.13 m ³ (STP) / h of oxygen molecules (purity greater than 99% by volume) were metered into the reaction gas A (T = 143 ° C, P = 2.40 bar). The metering addition was performed by throttling so that the produced pressure of the produced reaction gas A was still 1.81 bar.

The layer height of the second catalyst bed was such that the reaction gas A was spilled out in the second catalyst bed at a temperature of 520 DEG C and a pressure of 1.78 bar in the following amounts:

54.2% by volume of propane,

21.3% by volume of propylene,

1.90% by volume of methane,

0.63% by volume of ethylene,

H ₂ 3.61% by volume,

O ₂ 0 vol%, and

H ₂ O 17.3% by volume.

The effluent was 8.70 m ³ (STP) / h.

At the upstream of the stationary mixer located upstream of the third catalyst bed, 0.17 m ³ (STP) / h of oxygen molecules (above 99 vol% purity) were metered into this reaction gas A (T = 143 ° C, pressure 1.78 bar Lt; / RTI &gt; The pressure of the produced reaction gas A was still 1.78 bar.

The layer height of the third catalyst bed was such that the reaction gas A was spilled out as product gas A in the third catalyst bed at a temperature of 540 ° C and a pressure of 1.73 bar in the following amounts:

46.30% by volume of propane,

25.12% by volume of propylene,

0.75% by volume of ethylene,

2.27% by volume of methane,

4.50% by volume of H ₂ ,

O ₂ 0 vol%, and

H ₂ O 20.05% by volume.

The effluent was 9.17 m ³ (STP) / h.

The product gas A was divided into two parts of the same composition by a flow divider. The first part was 5.04 m ³ (STP) / h and the second part was 4.13 m ³ (STP) / h. The first portion was recycled to the reaction zone A as a component of the reaction gas A charge gas mixture. For the dehydrogenation circulating gas flow of the first part, the propane circulating gas heated as above is used as the driving jet to operate the jet pump, and the delivery direction of the driving jet compressed through the driving nozzle through the mixing zone and the diffuser is The suction nozzle indicates the direction of the first flow portion of the product gas A and the "suction nozzle-mixing zone-diffuser" connection refers to the first portion of the product gas A to be recycled and the second portion of the reaction zone A Thus forming a single connection line between the access points.

A second portion of the product gas A is delivered out of the reaction zone A and is first introduced into the reaction mixture A by a first indirect heat exchange with a first gas mixture comprised of crude propane and oxygen molecules, RTI ID = 0.0 &gt; 476 C. &lt; / RTI &gt; The working pressure of the product gas A was intrinsically maintained.

At the same time, the first gas mixture was heated to 23O &lt; 0 &gt; C to 450 &lt; 0 &gt; C and the working pressure of this gas mixture was essentially maintained at 1.84 bar.

In a second indirect indirect downstream heat exchange with the propane circulation gas from the separation zone IV, the product gas A was cooled from 476 캜 to 307 캜. Conversely, the temperature of the propane circulating gas increased from 69 ° C to 400 ° C.

Thereafter, steam of 0.52 m ³ (STP) / h was removed by condensation from product gas A cooled to 307 ° C. For this, the product gas A cooled to 307 ° C is first subjected to indirect heat exchange to produce product gas A ^* (T = 54 ° C, P = 1.73 bar) of 3.60 m ³ (STP) / h Removed and cooled to &lt; RTI ID = 0.0 &gt; 111 C). The temperature of the product gas A ^* increased to 277 ° C at the same time.

(29 [deg.] C) and directly removed from the product gas A by appropriate direct cooling, removes the aforementioned amount of water vapor from the product gas A by condensation to produce product gas A ^* (P = 1.73 bar, T = 54 캜, 3.60 m ³ (STP) / h). This was heated to 277 캜 as described above. Then, oxygen molecules of 1.39 m ³ (STP) / h (purity greater than 99% by volume, T = 143 ° C, P = 2.40 bar) were throttled with product gas A ^* . This formed a reaction gas B starting mixture (charging gas mixture) at a temperature of 262 ° C and a working pressure of 1.63 bar.

Reaction gas B The starting mixture (charge gas mixture) had the following contents:

38.26% by volume of propane,

20.77% by volume of propylene,

0.62% by volume of ethylene,

1.87% by volume of methane,

H ₂ 3.73% by volume,

27.53% by volume of O ₂ , and

H ₂ O 6.20% by volume.

This was used to charge reaction zone B having the same structure as in II.

During the first oxidation step the catalyst charge was chosen to be 164 l (STP) / l · h. The salt melt (53 wt.% KNO ₃ , 40 wt.% NaNO ₂ , 7 wt.% NaNO ₃ ) had the following inlet temperature:

T _A = 331 ° C T _B = 339 ° C

T _C = 275 ° C T _D = 282 ° C

And sufficient molecular oxygen (143 ℃, 2.40 bar), the first addition of the weighing in the product gas mixture of the oxidation step (throttling), the 2 O ₂ in the resulting charge gas mixture for the oxidation step: acrolein was a molar ratio of 0.72. The amount of acrolein loaded during the second oxidation step of the catalyst was 139 l (STP) / l · h. The inlet pressure to the second oxidation step was 1.57 bar. Reaction gas B was effluent at an intermediate cooler at a temperature of 260 占 폚 and the inflow temperature of the packed gas mixture to the second oxidation step was 256 占 폚.

The product gas mixture of the first oxidation step had the following contents:

17.51% by volume of acrolein,

1.18% by volume of acrylic acid,

38.12% by volume of propane,

1.03% by volume of propylene,

1.86% by volume of methane,

H ₂ 3.71% by volume,

0.62% by volume of ethylene,

O ₂ 4.18% by volume,

27.83% by volume of H ₂ O,

0.88% by volume of CO, and

2.07% by volume of CO ₂ .

Before entering the subsequent cooler, the temperature of the product gas in the first oxidation step was 335 ° C.

The product gas mixture (product gas B) in the second oxidation step had a temperature of 270 DEG C, a pressure of 1.55 bar and also had the following contents:

0.089% by volume of acrolein,

17.67% by volume of acrylic acid,

0.52% by volume of acetic acid,

37.99% by volume of propane,

1.02% by volume of propylene,

1.86% by volume of methane,

0.62% by volume of ethylene,

3.70% by volume of H ₂ ,

O ₂ 2.61% by volume,

28.42% by volume of H ₂ O,

1.41% by volume of CO, and

CO ₂ 3.10% by volume.

As described in WO 2004/035514, product gas B (5.03 m ³ (STP) / h) was fractionally condensed in a tray column (separation zone II).

As the first fuel, 13.9 g / h of high boiling point water (polyacrylic acid (Michael adduct), etc.) was fed to the residue painting.

2.82 kg / h of condensed crude acrylic acid having a temperature of 15 캜 and 96.99% by weight of acrylic acid was discharged from the second collecting tray to the tray column on the supply of the product gas B. The suspension crystals were separated from the mother liquor in a hydraulic scrubbing column and the mother liquor was recycled to the condensation column as described in WO 2004/035514, as described in WO 2004/035514, after a small amount of water was added and the suspension crystallized. The purity of the washed suspension crystals is greater than 99.87% by weight of acrylic acid and is suitable for the preparation of water- "superabsorbent" polymers for immediate hygiene applications.

The amount of acidic water condensate discharged into the condensing column from the third collection tray on the feed of product gas B but not recycled to the condensing column was 1.24 kg / h, the temperature was 33 ° C, and had the following contents:

0.73% by weight of acrolein,

4.95% by weight of acrylic acid,

5.16% by weight acetic acid, and

88.1% by weight of water.

At the top of the condensation column, 2.72 m ³ (STP) / h of residual gas was sparged in the separation zone II at a temperature of 33 ° C and a pressure of 1.20 bar in the following amounts:

0.03% by volume of acrolein,

0.02% by volume of acrylic acid,

0.01% by volume of acetic acid,

70.26% by volume of propane,

1.91% by volume of propylene,

1.15% by volume of ethylene,

6.93% by volume of H ₂ ,

3.44% by volume of methane,

O ₂ 4.83% by volume,

2.11% by volume of H ₂ O,

2.59% by volume of CO, and

5.73% by volume of CO2.

In the first compressor stage of the multistage radial compressor, the residual gas was compressed from 1.20 bar to 6.93 bar during which the temperature of the residual gas increased to 127 ° C.

(Hereinafter also referred to as "off-gas") in an indirect heat exchanger, which was cooled to 114 [deg.] C without forming condensate Lt; RTI ID = 0.0 &gt; 20 &lt; / RTI &gt; bar in the first expansion stage of the multistage expansion turbine. This heated off gas (0.69 m ³ (STP) / h). In the second expansion step, the off-gas was compressed to 1.10 bar, at which time the temperature was 25 [deg.] C, and then fed to the retentate.

In an indirect air cooler, the compressed residual gas (P = 6.93 bar and T = 114 占 폚) was cooled to 59 占 폚. In an additional compressor stage, it was compressed from 6.93 bar to 20 bar and simultaneously heated to 114 ° C. In a downstream indirect heat exchanger, it was cooled to 107 ° C without forming a condensate (the coolant was a "propane-scrubbed" residual gas, simultaneously heated to 30 ° C to 84 ° C). At the downstream of the additional indirect heat exchanger, the residual gas at 107 ° C was cooled to 35 ° C (coolant is surface water at approximately 25 ° C).

This condensed 44.8 g / h of water out of the residual gas. These aqueous condensates were likewise fed to the residue painting and sparged with other mentioned residues (supplying 5.95 m ³ (STP) / h of air). Steam (52 bar, 267 占 폚, 5.16 kg / h) was obtained by cooling the hot combustion gas and released around.

Residual gas (20 bar, 35 캜) still containing 2.67 m ³ (STP) / h of propane was fed to the lower compartment of the column with random packing (separation zone III).

At the top of this scrubbing column 20.39 kg / h of industrial tetradecane from Haltermann, Germany (as described in II) was introduced at an introduction temperature of 30 ° C (pressure 20 bar).

At the top of the scrubbing column, a "propane-scrubbed" residual gas (off-gas) was sparged at a pressure of 20 bar and a temperature of 30 ° C and an amount of 0.69 m ³ (STP)

1.38% by volume of propane,

0.04% by volume of propylene,

4.48% by volume of ethylene,

13.44% by volume of methane,

H ₂ 27.12% by volume,

O ₂ 18.89% by volume,

1.04% by volume of H ₂ O,

10.16% by volume of CO, and

CO ₂ 22.44% by volume.

This off-gas was compressed as described, indirect heat exchanged, and then either fed to the residue or drawn by flash.

An absorbent of 24.27 kg / h at a pressure of 20 bar and a temperature of 61 ° C was discharged from the bottom of the externally heated or uncooled "propane scrubbing column". It contained 15.63 wt% propane and 0.41 wt% propylene.

The adsorbate was heated to 69 [deg.] C in an indirect heat exchanger prior to delivery to the top of the downstream desorption column. The heat carrier used was a liquid effluent from the desorption column (20.39 kg / h), which was at a pressure of 3.26 bar at a temperature of 80 ° C and still contained 0.089% by weight dissolved propane. The mixture was further compressed to 20 bar by a pump, indirectly cooled to 30 ° C using surface water (25 ° C), and recycled to propane adsorption at the top of the "propane scrubbing column".

After heating to 69 [deg.] C, the absorbent is compressed to a pressure of 3.26 bar (for example by an inverted pump or valve) (in the case of an inverted pump the mechanical energy released is suitably also removed from the desorption column, The liquid effluent of the column), and the resulting two-phase mixture was transferred to the top of the desorption column.

A water vapor of 0.1 kg / h was fed to the desorption column (likewise a column with random packing) counter-current to the absorption from the top of the desorption column from the bottom as a stripping gas at a pressure of 5 bar and a temperature of 15 캜.

A heating coil having a temperature of 152 DEG C, a pressure of 5.00 bar, and an amount of 0.8 kg / h was similarly fed back to the water vapor rising the desorption column was placed in the desorption column.

At the top of the desorption column, 2.04 m ³ (STP) / h of propane circulating gas at a temperature of 69 ° C and a working pressure of 3.26 bar was spilled. Heated to 400 DEG C by indirect heat exchange with the product gas A at a temperature of 476 DEG C and then propane circulating gas is used as a driving jet to operate the jet pump to produce dehydrogenated circulating gas and propane circulating gas as a mixture starting from the reaction gas A And fed to fill reaction zone A with the mixture.

Propane circulating gas had the following contents:

Propane 93.07% by volume,

2.52% by volume of propylene, and

H ₂ O 3.40% by volume.

IV. A third exemplary method for producing acrylic acid from propane (described as normal operation)

The reaction zone A is a tray reactor as described under II. Inhomogeneously catalyzed partial propane dehydrogenation was carried out in the tray reactor without the use of a linear pass-through dehydrogenation circulating gas mode. The propane loading for all trays in all trays (calculated without inert materials) was 1500 l (STP) / l · h.

A reaction gas A charge gas mixture (starting mixture; T = 418 DEG C; P = 1.77 bar) of 3.80 m ³ (STP) / h was fed to the first catalyst bed in the flow direction, which had the following contents:

75.21% by volume of propane,

1.35% by volume of propylene,

0.01% by volume of methane,

H ₂ 8.60% by volume,

O ₂ 4.32% by volume,

0.96% by volume of N ₂ ,

7.68% by volume of H ₂ O,

0.02% by volume of CO, and

0.83% by volume of CO ₂ .

Reaction gas A The charge gas mixture is a mixture of a first gas mixture (1.59 m ³ (STP) / h) and a second gas mixture (2.21 m ³ (STP) / h) &Lt; / RTI &gt; The temperature of the first gas mixture was 450 DEG C and the pressure was 1.77 bar, consisting of:

1.02 m ³ (STP) / h crude propane (T = 237 ° C, P = 2.90 bar) containing propane (fresh propane) to above 98%

0.16 m ³ (STP) / h of O ₂ (T = 143 ° C, P = 1.77 bar, purity greater than 99% by volume), and

0.41 m ³ (STP) / h of gas (T = 30 ° C; throttle to 1.77 bar), which contains predominantly hydrogen molecules and is removed from the "propane-scrubbed" residual gas (off-gas) by membrane separation.

In the first heat exchanger, the indirect heat exchange with the product gas A delivered in the direction of the reaction zone B from outside the reaction zone A set the temperature of 450 ° C.

The second gas mixture was a propane circulating gas (2.21 m ³ (STP) / h; 400 ° C; 1.77 bar). Indirect heat exchange with the product gas A cooled to a temperature of 500 ° C in the first heat exchanger set the temperature of the propane circulating gas.

The inlet temperature of the Reaction Gas A packed gas mixture and the bed height of the first catalyst layer through which the Reaction Gas A packed gas mixture flows are such that the Reaction Gas A is spilled out in the catalyst bed at a temperature of 507 ° C and a pressure of 1.76 bar in the following amounts:

64.42% by volume of propane,

9.10% by volume of propylene,

0.079% by volume of ethylene,

0.33% by volume of methane,

H ₂ 7.63% by volume,

O ₂ 0% by volume,

0.92% by volume of N ₂ ,

15.69% by volume of H ₂ O,

0 vol% of CO, and

0.82% by volume of CO ₂ .

The runoff was 3.95 m ³ (STP) / h.

Over the first catalyst bed, 0.13 m ³ (STP) / h of oxygen molecules (purity greater than 99% by volume) were metered into the reaction gas A (T = 143 ° C, P = 2.40 bar). The metering addition was carried out by throttling so that the produced pressure of the produced reaction gas A was still 1.76 bar.

The layer height of the second catalyst layer was such that the reaction gas A was spilled out in the second catalyst layer at a temperature of 544 ° C and a pressure of 1.75 bar in the following amounts:

51.72% by volume of propane,

15.97% by volume of propylene,

0.149% by volume of ethylene,

0.60% by volume of methane,

8.8% by volume of H ₂ ,

O ₂ 0% by volume,

0.85% by volume of N ₂ ,

20.26% by volume of H ₂ O,

0 vol% of CO, and

0.76% by volume of CO ₂ .

The runoff was 4.29 m ³ (STP) / h.

An oxygen molecule (purity more than 99% by volume) of 0.12 m ³ (STP) / h was metered into this reaction gas A (T = 143 ° C, pressure of 1.78 bar) upstream of the stationary mixer located upstream of the third catalyst bed Lt; / RTI &gt; The pressure of the produced reaction gas A was still 1.75 bar.

The layer height of the third catalyst bed was such that reaction gas A was spilled out as product gas A in the third catalyst bed at a temperature of 571 ° C and a pressure of 1.73 bar in the following amounts:

40.12% by volume of propane,

22.2% by volume of propylene,

0.22% by volume of ethylene,

0.87% by volume of methane,

H ₂ 10.11% by volume,

O ₂ 0% by volume,

0.79% by volume of N ₂ ,

23.96% by volume of H ₂ O,

0 vol% of CO, and

0.70% by volume of CO ₂ .

The product gas A is delivered out of the reaction zone A and is first subjected to the first indirect heat exchange with the first gas mixture comprised in the reaction gas A starting mixture (charge gas mixture) and consisting of crude propane, hydrogen molecule- Lt; RTI ID = 0.0 &gt; 500 C &lt; / RTI &gt; The working pressure of the product gas A was intrinsically maintained. This heated the first gas mixture to 213 &lt; 0 &gt; C to 450 &lt; 0 &gt; C, and the working pressure of this gas mixture was maintained essentially 1.77 bar.

In the subsequent second indirect heat exchange with the propane circulating gas from the separation zone IV, the product gas A was cooled from 500 캜 to 337 캜. Conversely, the temperature of the propane circulating gas increased from 69 ° C to 400 ° C.

Thereafter, water vapor of 0.79 m ³ (STP) / h was separated from the product gas A cooled at 337 ° C by condensation. To this end, the product gas A cooled to 307 ° C is first subjected to indirect heat exchange to produce a product gas A ^* (T = 54 ° C, P = 1.73 bar) of 3.85 m ³ (STP) / h RTI ID = 0.0 &gt; 120 C). &Lt; / RTI &gt;

The temperature of the product gas A ^* increased to 307 ° C at the same time.

(29 [deg.] C) and directly removed from the product gas A by appropriate direct cooling, removes the aforementioned amount of water vapor from the product gas A by condensation to produce product gas A ^* (P = 1.73 bar, T = 54 캜, 3.85 m ³ (STP) / h). This was heated to 307 占 폚 as described above. Then, 1.41 m ³ (STP) / h of oxygen molecules (purity greater than 99% by volume, T = 143 ° C, P = 2.40 bar) were throttled with product gas A ^* . This formed a reaction gas B starting mixture (charging gas mixture) having a temperature of 290 ° C and a working pressure of 1.63 bar.

Reaction gas B The starting mixture (charge gas mixture) had the following contents:

35.35% by volume of propane,

19.60% by volume of propylene,

0.19% by volume of ethylene,

0.76% by volume of methane,

H ₂ 8.91% by volume,

O ₂ 26.60% by volume,

0.69% by volume of N ₂ ,

6.27% by volume of H ₂ O,

0 vol% of CO, and

0.62% by volume of CO ₂ .

This was used to charge reaction zone B having the same structure as in II. During the first oxidation step the catalyst charge was chosen to be 130 l (STP) / l · h. The salt melt (53 wt.% KNO ₃ , 40 wt.% NaNO ₂ , 7 wt.% NaNO ₃ ) had the following inlet temperature:

T _A = 322 ° C T _B = 327 ° C

T _C = 261 ° C T _D = 268 ° C

And sufficient molecular oxygen (143 ℃, 2.40 bar), the first addition of the weighing in the product gas mixture of the oxidation step (throttling), the 2 O ₂ in the resulting charge gas mixture for the oxidation step: acrolein molar ratio was 0.745. The amount of acrolein loaded during the second oxidation step of the catalyst was 110 l (STP) / l · h. The inlet pressure to the second oxidation step was 1.57 bar. Reaction gas B was effluent at an intermediate cooler at a temperature of 260 占 폚 and the inflow temperature of the packed gas mixture to the second oxidation step was 256 占 폚.

The product gas mixture of the first oxidation step had the following contents:

16.51% by volume of acrolein,

1.11% by volume of acrylic acid,

35.23% by volume of propane,

0.98% by volume of propylene,

0.19% by volume of ethylene,

0.76% by volume of methane,

H ₂ 8.88% by volume,

O ₂ 4.57% by volume,

0.69% by volume of N ₂ ,

26.66% by volume of H ₂ O,

0.83% by volume of CO, and

2.56% by volume of CO ₂ .

Before entering the subsequent cooler, the temperature of the product gas in the first oxidation step was 335 ° C.

The product gas mixture (product gas B) in the second oxidation step had a temperature of 270 DEG C, a pressure of 1.55 bar and also had the following contents:

0.08% by volume of acrolein,

16.71% by volume of acrylic acid,

0.495% by volume of acetic acid,

35.20% by volume of propane,

0.89% by volume of propylene,

0.19% by volume of ethylene,

0.76% by volume of methane,

H ₂ 8.87% by volume,

O ₂ 2.86% by volume,

0.69% by volume of N ₂ ,

27.28% by volume of H ₂ O,

1.33% by volume of CO, and

3.54% by volume of CO ₂ .

As described in WO 2004/035514, product gas B (5.293 m ³ (STP) / h) was fractionally condensed in a tray column (separation zone II).

As the first fuel, 13.8 g / h of high boiling point water (polyacrylic acid (Michael adduct), etc.) was supplied to the residue painting.

2.80 kg / h of condensed crude acrylic acid having a temperature of 15 캜 and 96.99% by weight of acrylic acid was discharged from the second collecting tray to the tray column on the supply of the product gas B. The suspension crystals were separated from the mother liquor in a hydraulic scrubbing column and the mother liquor was recycled to the condensation column as described in WO 2004/035514, as described in WO 2004/035514, after a small amount of water was added and the suspension crystallized. The purity of the washed suspension crystals is greater than 99.87% by weight of acrylic acid and is suitable for the preparation of water-"breathable" polymers for immediate hygiene applications.

The amount of acidic water condensate discharged into the condensation column from the third collection tray on the feed of product gas B but not recycled to the condensation column was 1.25 kg / h, the temperature was 33 ° C, and had the following contents:

0.72% by weight of acrolein,

4.95% by weight of acrylic acid,

5.09% by weight acetic acid, and

88.24% by weight of water.

At the top of the condensation column, 2.98 m ³ (STP) / h of residual gas was sparged in the separation zone II at a temperature of 33 ° C and a pressure of 1.20 bar in the following amounts:

0.02% by volume of acrolein,

0.02% by volume of acrylic acid,

62.39% by volume of propane,

1.73% by volume of propylene,

0.34% by volume of ethylene,

1.35% by volume of methane,

H ₂ 16.23% by volume,

O ₂ 5.05% by volume,

1.27% by volume of N ₂ ,

1.92% by volume of H ₂ O,

2.36% by volume of CO, and

6.33% by volume of CO ₂ .

In the first compressor stage of the multistage radial compressor, the residual gas was compressed from 1.20 bar to 6.93 bar during which the temperature of the residual gas increased to 134 ° C.

In a indirect heat exchanger, it was cooled to 123 ° C without formation of condensate (the coolant was "propane-scrubbed" residual gas (hereinafter also referred to as "off-gas"), And heated from 30 DEG C to 102 DEG C in an indirect heat exchanger, then compressed starting from 20 bar in the first expansion stage of the multistage expansion turbine and cooled at the same time). This heated off gas (0.59 m ³ (STP) / h). In the second expansion step, the off gas was compressed to 1.10 bar, at which time the temperature was 23 [deg.] C, and then fed to the retentate.

In an indirect air cooler, the compressed residual gas (P = 6.93 bar and T = 134 ° C) was cooled to 59 ° C. In an additional compressor stage, it was compressed from 6.93 bar to 20 bar and heated to 117 ° C at the same time. In a downstream indirect heat exchanger, it was cooled to 110 ° C without forming condensate (the coolant was "propane-scrubbed" residual gas at the completion of the H ₂ film removal and simultaneously heated from 30 ° C to 102 ° C). At the downstream of the additional indirect heat exchanger, the residual gas at 110 DEG C was cooled to 35 DEG C (the coolant is surface water at approximately 25 DEG C). This condensed 44.2 g / h of water out of the residual gas. This aqueous condensate was likewise fed to the residue painting and drawn with 4.12 m ³ (STP) / h of air in conjunction with the other residues mentioned. Steam (52 bar, 267 캜, 3.41 kg / h) was obtained by cooling the hot combustion gas and released to ambient.

Residual gas (20 bar, 35 캜) still containing 2.93 m ³ (STP) / h of propane was delivered to the lower compartment of the column with random packing (separation zone III). At the top of this scrubbing column, 21.62 kg / h of industrial tetradecane from Haltermann, Germany (as described in II) was introduced at an introduction temperature of 30 ° C (pressure 20 bar).

On the top of the scrubbing column, a "propane-scrubbed" residual gas (off-gas) was sparged at a pressure of 20 bar and a temperature of 30 ° C and a volume of 1.01 m ³ (STP)

0.92% by volume of propane,

0.03% by volume of propylene,

0.99% by volume of ethylene,

3.99% by volume of methane,

47.97% by volume of H ₂ ,

O ₂ 14.94% by volume,

N ₂ 3.75% by volume,

0.78% by volume of H ₂ O,

6.95% by volume of CO, and

CO ₂ 18.69% by volume.

This "propane-scrubbed" offgas was transferred to a bundle of cast tube membranes (external pressure: 1.77 bar; BH polyimide membrane from UBE Industries Ltd.) and the gas stream (infiltration stream) 0.41 m (P = 20 bar; T = 30 &lt; 0 &gt; C) with the release of ³ (STP) / h.

0.02% by volume of ethylene,

0.059% by volume of methane,

H ₂ 79.15% by volume,

O ₂ 1.53% by volume,

N ₂ 8.83% by volume,

1.52% by volume of H ₂ O,

0.22% by volume of CO, and

CO ₂ 7.68% by volume.

This gas stream was supplied (throttled) to the first gas mixture to produce a reactive gas A charge gas mixture.

The remaining off-gas was compressed and indirect heat exchanged as described, and then either fed to the residue filtration or by flashing.

An absorbent of 25.40 kg / h at a pressure of 20 bar and a temperature of 59 캜 was discharged from the bottom of the externally cooled or unheated "propane scrubbing column ". It contained 14.40 wt% propane and 0.38 wt% propylene.

The adsorbate was heated to 69 [deg.] C in an indirect heat exchanger prior to delivery to the top of the downstream desorption column. The heat carrier used was a liquid effluent from the desorption column (21.62 kg / h), which at a pressure of 3.26 bar was at a temperature of 80 ° C and still contained 0.09% by weight dissolved propane. The mixture was further compressed to 20 bar by a pump, indirectly cooled to 30 ° C using surface water (25 ° C), and recycled to propane adsorption at the top of the "propane scrubbing column".

After heating to 69 [deg.] C, the absorbent is compressed to a pressure of 3.26 bar (for example by an inverted pump or valve) (in the case of an inverted pump the mechanical energy released is suitably also removed from the desorption column, The liquid effluent of the column), and the resulting two-phase mixture was transferred to the top of the desorption column.

A water vapor of 0.2 kg / h was fed to the desorption column (likewise a column with random packing) counter-current to the absorption from the top of the desorption column from the bottom as a stripping gas at a pressure of 5 bar and a temperature of 152 캜.

A heating coil having a temperature of 152 DEG C, a pressure of 5.00 bar, and an amount of 0.8 kg / h was similarly fed back to the water vapor rising the desorption column was placed in the desorption column.

At the top of the desorption column, a propane circulation gas of 2.21 m ³ (STP) / h at a temperature of 69 ° C and a working pressure of 3.26 bar was drained. After being heated to 400 DEG C by indirect heat exchange with a product gas A at a temperature of 476 DEG C, a propane circulating gas was fed to fill the reaction zone A with the reaction mixture A as a starting mixture.

Propane circulating gas had the following contents:

83.73% by volume of propane,

2.32% by volume of propylene, and

H ₂ O 12.94% by volume.

V. Fourth example process for producing acrylic acid from propane (described as normal operation)

Reaction zone A is a tray reactor as described under II above except that its connected downstream in this embodiment has only one fixed catalyst bed and its arrangement is in the direction of flow from the tray reactor to an additional heat insulating shaft corresponding to the first fixed bed tray Except that it is a reactor. An indirect heat exchanger is connected between the two reactors. Inhomogeneously catalyzed partial propane dehydrogenation was carried out in the tray reactor using a linear pass-through dehydrogenation recycle gas mode. The propane loading for all trays in all trays (calculated without inert materials) was 1500 l (STP) / l · h. In the reactor as a downstream shaft, essentially only heterogeneous catalytic combustion of hydrogen molecules was performed.

A feed gas mixture (starting mixture; P = 2.46 bar) of 6.93 m ³ (STP) / h was fed to the first catalyst bed of the tray reactor in the flow direction, which had the following contents:

60.54% by volume of propane,

10.38% by volume of propylene,

0.28% by volume of ethylene,

0.87% by volume of methane,

H ₂ 2.44% by volume,

O ₂ 1.19% by volume,

N ₂ 16.48% by volume,

6.82% by volume of H ₂ O,

0 vol% of CO, and

0.01% by volume of CO ₂ .

Reaction gas A The charge gas mixture is a mixture of a first gas mixture (1.47 m ³ (STP) / h) and a second gas mixture (5.46 m ³ (STP) / h) &Lt; / RTI &gt;

The temperature of the first gas mixture was 500 DEG C and the pressure was 2.46 bar, consisting of:

1.06 m ³ (STP) / h crude propane (T = 237 ° C, P = 2.90 bar) containing propane (fresh propane) to an extent of greater than 98%

0.41 m ³ (STP) / h of compressed air (T = 159 ° C; P = 2.66 bar).

Indirect heat exchange (in the intermediate cooler, heat exchanger disposed between the two reactors) with the reaction gas A delivered from the outside of the tray reactor to the downstream shaft reactor was set at a temperature of 500 ° C.

The second gas mixture was composed of:

3.48 m ³ (STP) / h of dehydrogenation circulating gas (T = 554 ° C; P = 2.37 bar), and

Propane circulating gas of 1.98 m ³ (STP) / h (T = 500 ° C; P = 2.97 bar).

Its pressure was 2.46 bar and the temperature was 529 ° C. Indirect heat exchange with the product gas A delivered outside the downstream shaft reactor set the temperature of the propane circulating gas. Dehydrogenation The circulating gas flow was carried out in accordance with the jet pump principle (DE-A 10211275) and the propane circulating gas functioning as the drive jet was heated to 500 ° C.

The inlet temperature of the reaction gas A packed gas mixture and the layer height of the first catalyst layer in which the reaction gas A packed gas mixture flows in the flow direction in the tray reactor are such that the reaction gas A flows out from the catalyst layer at a temperature of 522 캜 and a pressure of 2.44 bar :

53.77% by volume of propane,

14.45% by volume of propylene,

0.29% by volume of ethylene,

1.26% by volume of methane,

H ₂ 4.31% by volume,

O ₂ 0% by volume,

N ₂ 15.91% by volume,

8.8% by volume of H ₂ O,

0 vol% of CO, and

0.01% by volume of CO ₂ .

The effluent was 7.18 m ³ (STP) / h. 0.64 m ³ (STP) / h of air was metered into the reaction gas A (T = 159 ° C, P = 2.66 bar) through the first catalyst layer of the tray reactor. The metering addition was carried out by throttling so that the produced pressure of the produced reaction gas A was still 2.44 bar. The layer height of the second catalyst layer of the tray reactor was such that the reaction gas A was spilled out in the second catalyst layer at a temperature of 539 ° C and a pressure of 2.41 bar in the following amounts:

43.71% by volume of propane,

17.02% by volume of propylene,

0.51% by volume of ethylene,

1.52% by volume of methane,

4.53% by volume of H ₂ ,

O ₂ 0% by volume,

20.35% by volume of N ₂ ,

11.37% by volume of H ₂ O,

0 vol% of CO, and

0.01% by volume of CO ₂ .

The effluent was 8.04 m ³ (STP) / h.

0.63 m ³ (STP) / h of air was metered into this reaction gas A (T = 159 ° C, throttled at a pressure of 2.66 bar) upstream of the stationary mixer located upstream of the third catalyst bed of the tray reactor. The pressure of the produced reaction gas A was still 2.41 bar.

The layer height of the third catalyst layer of the tray reactor was such that the reactive gas A was spilled in the third catalyst layer at a temperature of 554 DEG C and a pressure of 2.37 bar in the following amounts:

35.58% by volume of propane,

19.15% by volume of propylene,

0.57% by volume of ethylene,

1.72% by volume of methane,

H ₂ 4.82% by volume,

O ₂ 0% by volume,

N ₂ 23.83% by volume,

13.31% by volume of H ₂ O,

0 vol% of CO, and

0.01% by volume of CO ₂ .

The effluent was 8.89 m ³ (STP) / h.

At the exit of the tray reactor with a flow divider, the reaction gas A was divided into two portions of the same composition. The first part was 3.49 m ³ (STP) / h and the second part was 5.40 m ³ (STP) / h. The first portion was recycled to the reaction zone A as a component of the reaction gas A charge gas mixture. For the dehydrogenation circulating gas flow of the first part, the propane circulating gas heated as above is used as the driving jet to operate the jet pump, and the delivery direction of the driving jet compressed through the driving nozzle through the mixing zone and the diffuser is The suction nozzle indicates the direction of the first flow portion of the reaction gas A and the "suction nozzle-mixing zone-diffuser" connection indicates the direction of the flow of the reaction gas A into the first part of the reaction gas A to be recycled, Thus forming a single connection line between the access points.

Transferring a second portion of the reaction gas A out of the reaction zone A and first of all comprising the reaction gas A in the intermediate mixture (charge gas mixture) in the intermediate cooler disposed between the tray reactor and its shaft reactor downstream and consisting of crude propane and oxygen molecules Lt; RTI ID = 0.0 &gt; 554 C &lt; / RTI &gt; by first indirect heat exchange with the first gas mixture. This heated the first gas mixture to 22O &lt; 0 &gt; C to 500 &lt; 0 &gt; C, and the working pressure was essentially maintained.

Then, 6406 m ³ (STP) / h of air was metered into the second portion of the cooled reaction gas A as described above (T = 159 ° C, throttled to 2.66 bar). The gas mixture was then passed through a downstream shaft reactor. The layer height of the ("fourth") catalyst layer disposed in the downstream shaft reactor was such that reaction gas A was spilled into product gas A in the catalyst layer at a temperature of 580 ° C. and a pressure of 2.26 bar in the following amounts:

32.53% by volume of propane,

17.48% by volume of propylene,

0.52% by volume of ethylene,

1.57% by volume of methane,

H ₂ 0 vol%

O ₂ 0% by volume,

N ₂ 30.08% by volume,

16.79% by volume of H ₂ O,

0 vol% of CO, and

0.01% by volume of CO ₂ .

In downstream (second) indirect heat exchange with propane circulating gas from separation zone IV, product gas A (5.92 m ³ (STP) / h) was cooled from 580 ° C to 388 ° C. Conversely, the temperature of the propane circulating gas increased from 69 ° C to 500 ° C.

Thereafter, water vapor of 0.82 m ³ (STP) / h was removed by condensation from product gas A cooled to 388 ° C. To this end, the product gas A cooled to 388 ° C is first subjected to indirect heat exchange to produce a product gas A ^* (T = 40 ° C; P = 2.01 bar) of 5.10 m ³ (STP) Removed and cooled to 112 [deg.] C). In the downstream air cooler, the product gas A (T = 112 DEG C; P = 2.26 bar) was cooled to 54 DEG C by indirect heat exchange with outside air (the working pressure of product gas A dropped simultaneously to 2.06 bar) This was followed by two phases.

(29 [deg.] C) and directly removed from the product gas A by appropriate direct cooling, removes the aforementioned amount of water vapor from the product gas A by condensation to produce product gas A ^* (P = 2.01 bar, T = 40 캜, 5.10 m ³ (STP) / h). This was heated to 358 [deg.] C as described above. Subsequently, 8.55 m ³ (STP) / h of air (T = 159 ° C, P = 2.66 bar) was throttled with product gas A ^* . This formed a reaction gas B starting mixture (charging gas mixture) at a temperature of 283 DEG C and a working pressure of 1.96 bar.

Reaction gas B The starting mixture (charge gas mixture) had the following contents:

14.11% by volume of propane,

7.58% by volume of propylene,

0.23% by volume of ethylene,

0.68% by volume of methane,

H ₂ 0 vol%

O ₂ 12.69% by volume,

N ₂ 60.9% by volume,

2.75% by volume of H ₂ O,

0 vol% of CO, and

0.03% by volume of CO ₂ .

This was used to charge reaction zone B having the same structure as in II.

During the first oxidation step the catalyst charge was chosen to be 130 l (STP) / l · h. The salt melt (53 wt.% KNO ₃ , 40 wt.% NaNO ₂ , 7 wt.% NaNO ₃ ) had the following inlet temperature:

T _A = 327 ° C T _B = 329 ° C

T _C = 267 ° C T _D = 268 ° C

Sufficient air (159 캜, 2.66 bar) was metered into the product gas mixture of the first oxidation step (throttling), so that the molar ratio of O ₂ : acrolein in the resulting packed gas mixture for the second oxidation step was 1.01. The amount of acrolein loaded during the second oxidation step of the catalyst was 110 l (STP) / l · h. The inlet pressure to the second oxidation step was 1.68 bar. Reaction gas B exited the intermediate cooler at a temperature of 260 占 폚, and the inlet temperature of the charge gas mixture to the second oxidation step was 253 占 폚.

The product gas mixture of the first oxidation step had the following contents:

6.40% by volume of acrolein,

0.44% by volume of acrylic acid,

14.9% by volume of propane,

0.39% by volume of propylene,

0.68% by volume of methane,

H ₂ 0 vol%

0.23% by volume of ethylene,

O ₂ 4.18% by volume,

N ₂ 60.87% by volume,

H ₂ O 10.64% by volume,

0.33% by volume of CO, and

0.78% by volume of CO ₂ .

Before entering the subsequent cooler, the temperature of the product gas in the first oxidation step was 335 ° C.

The product gas mixture (product gas B) in the second oxidation step had a temperature of 270 DEG C, a pressure of 1.55 bar and also had the following contents:

0.03% by volume of acrolein,

5.97% by volume of acrylic acid,

0.18% by volume of acetic acid,

13.0% by volume of propane,

0.36% by volume of propylene,

0.63% by volume of methane,

0.21% by volume of ethylene,

H ₂ 0 vol%

O ₂ 2.60% by volume,

N ₂ 64.2% by volume,

H ₂ O 10.29% by volume,

0.48% by volume of CO, and

CO ₂ 1.08% by volume.

As described in WO 2004/035514, product gas B (14.81 m ³ (STP) / h) was fractionally condensed in a tray column (separation zone II).

As the first fuel, 13.8 g / h of high boiling point water (polyacrylic acid (Michael adduct), etc.) was supplied to the residue painting.

2.80 kg / h of condensed crude acrylic acid having a temperature of 15 캜 and 96.99% by weight of acrylic acid was discharged from the second collecting tray to the tray column on the supply of the product gas B. The suspension crystals were separated from the mother liquor in a hydraulic scrubbing column and the mother liquor was recycled to the condensation column as described in WO 2004/035514, as described in WO 2004/035514, after a small amount of water was added and the suspension crystallized. The purity of the washed suspension crystals is greater than 99.87% by weight of acrylic acid and is suitable for the preparation of water- "superabsorbent" polymers for immediate hygiene applications.

The amount of acidic water condensate discharged to the condensing column from the third collection tray on the feed of product gas B but not recycled to the condensing column was 1.19 kg / h, the temperature was 34 ° C, and had the following contents:

0.35% by weight of acrolein,

4.95% by weight of acrylic acid,

5.31% by weight acetic acid, and

Water 88.40% by weight.

At the top of the condensation column, a residual gas of 12.56 m ³ (STP) / h was sparged in the separation zone II at a temperature of 33 ° C and a pressure of 1.20 bar in the following amounts:

0.02% by volume of acrolein,

0.02% by volume of acrylic acid,

0% by volume of acetic acid,

15.35% by volume of propane,

0.42% by volume of propylene,

0.25% by volume of ethylene,

0.74% by volume of methane,

H ₂ 0 vol%

O ₂ 3.07% by volume,

N ₂ 75.64% by volume,

1.68% by volume of H ₂ O,

0.55% by volume of CO, and

1.27% by volume of CO ₂ .

In the first compressor stage of the multistage radial compressor, the residual gas was compressed from 1.20 bar to 6.93 bar during which the temperature of the residual gas increased to 217 ° C.

(Hereinafter also referred to as "off-gas") in an indirect heat exchanger without cooling to a temperature of 111 [deg.] C Lt; RTI ID = 0.0 &gt; 20 &lt; / RTI &gt; bar in the first expansion stage of the multistage expansion turbine and cooled at the same time). This heated off gas (10.39 m ³ (STP) / h). In the second expansion step, the off gas was compressed to 1.10 bar, at which time the temperature was 23 [deg.] C, and then fed to the retentate.

In an indirect air cooler, the compressed residual gas (P = 6.93 bar and T = 111 ° C) was cooled to 59 ° C. In an additional compressor stage, it was compressed from 6.93 bar to 20 bar and simultaneously heated to 168 ° C. In a downstream indirect heat exchanger, it was cooled to 101 ° C without forming condensate (the coolant was a "propane-scrubbed" residual gas and was simultaneously heated to 30 ° C to 138 ° C). Downstream of the additional indirect heat exchanger, the residual gas at 101 DEG C was cooled to 35 DEG C (the coolant is surface water at approximately 25 DEG C). This condensed 159.7159.7 g / h of water outside the residual gas. This aqueous condensate was likewise fed to the residue painting and drawn with 5.59 m ³ (STP) / h of air with other mentioned residues. Steam (52 bar, 267 캜, 3.78 kg / h) was obtained by cooling the hot combustion gas and released to ambient.

Residual gas (20 bar, 35 캜) still containing 12.37 m ³ (STP) / h of propane was delivered to the lower compartment of the column with random packing (separation zone III). At the top of this scrubbing column, 66.37 kg / h of industrial tetradecane from Haltermann, Germany (as described in II above) was introduced at an introduction temperature of 30 ° C (pressure 20 bar).

On the top of the scrubbing column, a "propane-scrubbed" residual gas (off-gas) was sparged at a pressure of 20 bar and a temperature of 30 ° C and a volume of 10.39 m ³ (STP)

0.19% by volume of propane,

0.01% by volume of propylene,

0.29% by volume of ethylene,

0.89% by volume of methane,

H ₂ 0 vol%

O ₂ 3.70% by volume,

N ₂ 91.38% by volume,

0.32% by volume of H ₂ O,

0.67% by volume of CO, and

1.53% by volume of CO ₂ .

This off-gas was compressed as described, indirect heat exchanged, and then either fed to the residue or drawn by flash.

An absorption of 70.25 kg / h of a pressure of 20 bar and a temperature of 41 ° C was discharged from the bottom of the externally cooled or unheated "propane scrubbing column". This included 5.41 wt% propane and 0.14 wt% propylene. The adsorbate was heated to 69 [deg.] C in an indirect heat exchanger prior to delivery to the top of the downstream desorption column. The heat carrier used was a liquid effluent from the desorption column (66.37 kg / h), which at a pressure of 2.97 bar was at a temperature of 80 ° C and still contained 0.03% by weight dissolved propane. The mixture was further compressed to 20 bar by a pump, indirectly cooled to 30 ° C using surface water (25 ° C), and recycled to propane adsorption at the top of the "propane scrubbing column".

After heating to 69 占 폚, the absorbent is compressed to a pressure of 2.97 bar (for example by an inverted pump or valve) (in the case of an inverted pump the mechanical energy released is suitably also removed from the desorption column, The liquid effluent of the column), and the resulting two-phase mixture was transferred to the top of the desorption column.

A heating coil with a temperature of 152 DEG C, a pressure of 5.00 bar, and a quantity of 1.3 kg / h was fed in parallel to the absorber derived from the desorption column was placed in the desorption column.

At the top of the desorption column, 1.98 m ³ (STP) / h of propane circulating gas at a temperature of 69 ° C and a working pressure of 2.97 bar was drained. After heating to 500 DEG C by indirect heat exchange with a product gas A at a temperature of 580 DEG C, the propane circulating gas was used as a driving jet to operate the jet pump, and the dehydrogenated circulating gas and the propane circulating gas were fed into the reaction mixture A The reaction zone A was fed to fill.

Propane circulating gas had the following contents:

96.42% by volume of propane, and

2.58% by volume of propylene.

VI. Comparative Example 1 of the method for producing acrylic acid from propane (described as normal operation)

The reaction zone A consisted of a reactor with an insulating shaft with only one fixed bed and an arrangement corresponding to the first fixed bed tray in the direction of flow in the tray reactor of II.

Inhomogeneously catalyzed partial propane dehydrogenation was carried out in the shaft reactor using a linear pass-through dehydrogenation recycle gas mode. The propane loading for the total catalyst (calculated without inert material) was 1500 l (STP) / l · h.

Reacting gas A 24.87 m ³ (STP) / h The charge gas mixture (starting mixture; T = 450 ° C; P = 2.82 bar) was fed to the reactor with a shaft, which had the following contents:

22.3% by volume of propane,

3.65% by volume of propylene,

0.07% by volume of ethane,

0.17% by volume of methane,

H ₂ 4.18% by volume,

O ₂ 2.09% by volume,

N ₂ 60.32% by volume,

5.05% by volume of H ₂ O,

0 vol% of CO, and

CO ₂ 1.17% by volume.

Reaction gas A The charge gas mixture is a mixture of a first gas mixture (1.69 m ³ (STP) / h) and a second gas mixture (23.18 m ³ (STP) / h) &Lt; / RTI &gt;

The first gas mixture had a vapor pressure of 0.013 m ³ (STP) / h of water vapor (T = 134 ° C; P = 3 bar), 1.06 m ³ (STP) / h of propane (new propane) A mixture of hydrogen molecules (purity greater than 99% by volume; T = 35 캜; P = 2.20 bar) of crude propane (T = 122 ° C; P = 3.50 bar) and 0.62 m ³ (STP) (T = 444 ° C; P = 2.82 bar) of 11.41 m ³ (STP) / h of dehydrogenated circulating gas (T = 534 ° C; P = 2.36 bar) and 11.16 m ³ ) And 0.61 m ³ (STP) / h of compressed air (T = 175 ° C, P = 3.00 bar).

The indirect heat exchange of the product gas A delivered in the direction of the reaction zone B from outside the reaction zone A set the compressed residual gas circulation gas temperature 444 ° C.

Dehydrogenation The circulating gas flow was carried out in accordance with the jet pump principle (DE-A 102 112 75) and the compressed residual gas circulating gas functioning as the drive jet was set at 444 ° C. The compressed air was metered into the resulting mixture (throttling).

The layer height of the fixed catalyst bed through which the reactive gas A packed gas mixture flows is such that product gas A spills out from the catalyst bed at a temperature of 534 ° C and a pressure of 2.36 bar in the following amounts:

17.83% by volume of propane,

7.51% by volume of propylene,

0.15% by volume of ethane,

0.37% by volume of methane,

H ₂ 3.76% by volume,

O ₂ 0% by volume,

N ₂ 59.18% by volume,

9.06% by volume of H ₂ O,

0 vol% of CO, and

CO ₂ 1.14% by volume.

The runoff was 25.35 m ³ (STP) / h.

The flow divider divides the product gas A at the outlet of the shaft reactor into two parts of the same composition.

The first part was 11.41 m ³ (STP) / h and the second part was 13.94 m ³ (STP) / h. The first portion was recycled to the reaction zone A as a component of the reaction gas A charge gas mixture. For the dehydrogenation circulating gas flow of the first part, the jet pump is operated using the heated compressed residual gas circulating gas as the driving jet, and the delivery of the driving jet compressed through the driving nozzle through the mixing zone and the diffuser Direction refers to the direction of entry into the reaction zone A, the suction nozzles indicate the direction of the first flow portion of the product gas A, and the "suction nozzle-mixing zone-diffuser" connection comprises a first part of the product gas A to be recycled, A &lt; / RTI &gt;

The second part of the product gas A is delivered out of the reaction zone A and indirectly with the compressed residual gas circulating gas (11.16 m ³ (STP) / h; T = 131 ° C; P = 3.00 bar) in the indirect heat exchanger disposed downstream Cooled to 317 캜 by heat exchange (the working pressure is essentially maintained). This heated the residual gas circulation gas to 444 ° C and reduced the working pressure to 2.82 bar.

The product gas A was then cooled to 221 ° C in a second indirect heat exchange with "propane-scrubbed" product gas A (9.06 m ³ (STP) / h; T = 30 ° C, P = 10.50 bar) At the same time, the work pressure was reduced to 2.18 bar.

This heated "propane-scrubbed" product gas A (hereinafter also referred to as "off-gas") was heated to 287 ° C and the working pressure was changed to 10.10 bar. The offgas was then compressed from 10.10 bar to 3.50 bar in the first expansion stage of the multistage expansion turbine, and at the same time cooled from 287 ° C to 176 ° C. Then, the off-gas out of the first expansion step was transferred to the product gas A at 221 캜 in indirect heat exchange. The latter was essentially cooled to 215 ° C while maintaining the working pressure. This heated the off-gas to 191 ° C, at which time the working pressure was substantially maintained. The off-gas was then compressed to 1.23 bar in the second expansion stage of the multistage expansion turbine and cooled to 89 占 폚. Subsequently, off-gas was supplied to the residue painting.

The product gas A at 215 ° C and 2.18 bar was then first cooled to 60 ° C in an air-cooled indirect heat exchanger and then cooled to 39 ° C in a chilled water-cooled indirect heat exchanger. The condensate removed during the cooling process was removed by a liquid separator (0.68 kg / h).

The remaining 13.09 m ³ (STP) / h of product gas A was compressed to 4.78 bar in the first compressor stage of the multistage radial compressor. This heated the product gas A at 39 占 폚 to 106 占 폚. Thereafter, the product gas A was cooled to 54 DEG C by air in an indirect heat exchanger, and the condensed water (8.0 g / h) was removed by a liquid separator. In the second compression stage, the remaining 13.08 m ³ (STP) / h of product gas A was compressed to 10.50 bar. This is associated with the heating of the product gas A to 123 [deg.] C. The product gas A was first cooled to 54 DEG C by air cooling in an indirect heat exchanger and then cooled to 29 DEG C by indirect heat exchange with cooling water. The condensate removed during the cooling process was removed by a liquid separator (0.29 kg / h). The remaining product gas A of 12.71 m ³ (STP) / h (the working pressure is still essentially 10.50 bar) was transferred to the lower part of the column with random packing. Its content was as follows:

19.55% by volume of propane,

8.24% by volume of propylene,

0.16% by volume of ethane,

0.41% by volume of methane,

H ₂ 4.13% by volume,

O ₂ 0% by volume,

N ₂ 64.9% by volume,

0.38% by volume of H ₂ O,

0 vol% of CO, and

CO ₂ 1.25% by volume.

At the top of the propane scrubbing column 120.32 kg / h of industrial tetradecane from Haltermann, Germany (as described in II) was introduced at an introduction temperature of 30 ° C (pressure 10.50 bar).

At the top of the scrubbing column, "propane-scrubbed" product gas A (off-gas) was sparged at a pressure of 10.50 bar and a temperature of 30 ° C and a volume of 9.06 m ³ (STP)

0.02% by volume of propane,

0.01% by volume of propylene,

0.14% by volume of ethane,

0.56% by volume of methane,

H ₂ 5.79% by volume,

O ₂ 0% by volume,

N ₂ 91.0% by volume,

0.40% by volume of H ₂ O,

0 vol% of CO, and

CO ₂ 1.07% by volume.

After indirect heat exchange with the product gas A delivered in the reaction zone A as described above and multistage compression to 1.23 bar in a multistage expansion turbine, off-gas was either supplied to the residue painting or flash-coded. It may be appropriate to remove the hydrogen molecules present on the industrial scale before compressing offgas. This can be done, for example, by applying a membrane separation as described above to off-gas. As a result, the removed hydrogen molecules can be completely or partially recycled to the reaction zone A (suitably as a component of the reaction gas A fill gas mixture). Alternatively, for example, it can be combusted in a fuel cell.

127.37 kg / h of absorbent at a pressure of 10.5 bar and a temperature of 40 占 폚 was discharged from the bottom of the exothermally cooled or unheated "propane scrubbing column". It contained 3.83 wt% propane, 1.54 wt% propylene, 0.1 wt% ethane, and 0.1 wt% CO ₂ . The absorber was compressed to a pressure of 2.01 bar before it was delivered to the top of the downstream desorption column (for example with an inverted pump or valve; the mechanical energy released in the case of the retro pump was suitably also removed from the desorption column, (Used to recompress the liquid effluent of the desorption column). The resulting two-phase mixture was transferred to the top of the desorption column.

8.00 m ³ (STP) / h compressed to a pressure of 3.00 bar (heated to 175 ° C) was added to the desorption column (likewise with random packing) back to the absorber from the top of the desorption column, .

The liquid effluent (120.32 kg / h) of the desorption column was again pumped to 10.50 bar by a pump and indirectly cooled with cooling water and then recycled to propane adsorption at the top of the propane scrubbing column (adsorption column) at a temperature of 30 ° C.

At the top of the desorption column, a gas mixture of 11.64 m ³ (STP) / h was sparged at a temperature of 33 ° C and a pressure of 2.01 bar in the following amounts:

21.33% by volume of propane,

8.99% by volume of propylene,

0.07% by volume of ethane,

O ₂ 13.93% by volume,

N ₂ 52.6% by volume,

1.53% by volume of H ₂ O,

0 vol% of CO, and

0.55% by volume of CO ₂ .

Indirect heat exchange with water vapor increased the temperature of the gas mixture to 140 ° C (which reduced the working pressure to 1.91 bar).

This gas mixture was used as the reaction gas B starting mixture (charging gas mixture) to charge the reaction zone B having the same structure as the above II.

The propylene loading of the catalyst charge in the first oxidation step was selected to be 130 l (STP) / l · h. The salt melt (53 wt.% KNO ₃ , 40 wt.% NaNO ₂ , 7 wt.% NaNO ₃ ) had the following inlet temperature:

T _A = 329 ° C T _B = 330 ° C

T _C = 265 ° C T _D = 266 ° C

Reaction gas B having a temperature of 260 DEG C, a pressure of 1.68 bar and a content of 11.66 m &lt; ³ &gt; (STP) / h was spilled out from the intercooler:

7.51% by volume of acrolein,

0.51% by volume of acrylic acid,

21.3% by volume of propane,

0.53% by volume of propylene,

0.07% by volume of ethane,

O &lt; _{2 &gt;} 3.92 vol%

52.52% by volume of N ₂ ,

H ₂ O 10.81% by volume,

0.38% by volume of CO, and

CO ₂ 1.43% by volume.

(Throttling) 2.17 m ³ (STP) / h of air compressed to 3 bar (heated to 175 ° C) in the radial compressor so that the reaction gas inlet temperature to the second oxidation stage is 252 at inlet pressure 1.68 bar (Acrolein loading = 110 l (STP) / l · h).

The product gas mixture (product gas B) in the second oxidation step had a temperature of 270 DEG C and a pressure of 1.55 bar and had the following contents:

0.03% by volume of acrolein,

6.62% by volume of acrylic acid,

0.2% by volume of acetic acid,

18.51% by volume of propane,

0.46% by volume of propylene,

0.06% by volume of ethane,

O ₂ 2.95% by volume,

N ₂ 58.0% by volume,

10.0% by volume of H ₂ O,

0.52% by volume of CO, and

CO ₂ 1.64% by volume.

The product gas B (13.42 m ³ (STP) / h) was fractionally condensed in a tray column as described in WO 2004/035514.

High boiling point water (polyacrylic acid (Michael adduct) etc.) of 13.8 g / h was supplied to the residue painting.

2.82 kg / h of condensed crude acrylic acid having a temperature of 15 캜 and 96.99% by weight of acrylic acid was discharged from the second collecting tray to the tray column on the supply of the product gas B. After addition of a small amount of water, it was subjected to suspension-crystallization as described in WO 2004/035514, the suspension crystals were separated from the mother liquor in a hydraulic scrubbing column and the mother liquor was recycled to the condensation column as described in WO 2004/035514. The purity of the washed suspension crystals is greater than 99.87% by weight of acrylic acid and is suitable for the preparation of water- "superabsorbent" polymers for immediate hygiene applications.

The amount of acidic water condensate discharged into the condensation column from the third collection tray on the feed of product gas B but not recycled to the condensation column was 1.04 kg / h and the temperature was 33 ° C. Likewise, the residue was supplied to the painting.

At the top of the condensation column, a residual gas of 11.16 m ³ (STP) / h was sparged in a condensation column at a temperature of 33 ° C and a pressure of 1.20 bar in the following amounts:

22.06% by volume of propane,

0.46% by volume of propylene,

3.54% by volume of O ₂ ,

N ₂ 69.75% by volume,

1.76% by volume of H ₂ O,

0 vol% of CO, and

CO ₂ 1.43% by volume.

The entirety of the residual gas is compressed to 3 bar (heated to 131 캜), indirectly exchanged with the product gas A delivered in the reaction zone A as described, (the drive for the jet pump used in the dehydrogenation recycle gas flow Lt; RTI ID = 0.0 &gt; A) &lt; / RTI &gt;

VII. Comparative Example 2 of the method for producing acrylic acid from propane (described as a normal working state)

The reaction zone A consisted of a reactor with an adiabatic shaft having only one fixed catalyst bed and an arrangement corresponding to the first fixed bed tray in the flow direction in the tray reactor of II.

Inhomogeneously catalyzed partial propane dehydrogenation was carried out in the shaft reactor using a linear pass-through dehydrogenation recycle gas mode. The propane loading for the total catalyst (calculated without inert material) was 1500 l (STP) / l · h.

Reaction gas A charge gas mixture (starting mixture; T = 456 ° C; P = 2.82 bar) of 23.79 m ³ (STP) / h was fed to the reactor with a shaft, which had the following contents:

19.31% by volume of propane,

3.82% by volume of propylene,

0.11% by volume of ethane,

0.33% by volume of methane,

H ₂ 4.27% by volume,

O ₂ 2.14% by volume,

62.72% by volume of N ₂ ,

H ₂ O 5.11% by volume,

0 vol% of CO, and

CO ₂ 1.22% by volume.

Reaction gas A The charge gas mixture is a mixture of a first gas mixture (1.69 m ³ (STP) / h) and a second gas mixture (22.10 m ³ (STP) / h) &Lt; / RTI &gt;

The first gas mixture had a vapor pressure of 0.013 m ³ (STP) / h of water vapor (T = 134 ° C; P = 3 bar), 1.08 m ³ (STP) / h of propane (new propane) (Purity greater than 99% by volume; T = 35 캜; P = 2.20 bar) of crude propane (T = 122 ° C; P = 3.50 bar) and 0.60 m ³ (STP) (P = 3 bar; T = 455 ° C) at 10.93 m ³ (STP) / h of dehydrogenation circulating gas (T = 546 ° C; P = 2.82 bar) and 10.52 m ³ ) And 0.65 m ³ (STP) / h of compressed air (T = 175 ° C, P = 3.00 bar).

The indirect heat exchange of the product gas A delivered in the direction of the reaction zone B outside the reaction zone A set the temperature of the compressed residual gas circulation gas at 455 ° C.

Dehydrogenation The circulating gas flow was carried out in accordance with the jet pump principle (DE-A 102 112 75) and the compressed residual gas circulating gas functioning as the drive jet was set at 455 ° C. The compressed air was metered into the resulting mixture (throttling).

The bed height of the fixed catalyst bed in which the reactive gas A packed gas mixture flows is such that the product gas A flows out in the catalyst bed at a temperature of 546 ° C and a pressure of 2.39 bar in the following amounts:

14.60% by volume of propane,

7.84% by volume of propylene,

0.24% by volume of ethane,

0.70% by volume of methane,

H ₂ 3.79% by volume,

O ₂ 0% by volume,

N ₂ 61.44% by volume,

9.19% by volume of H ₂ O,

0 vol% of CO, and

CO ₂ 1.19% by volume.

The runoff was 24.29 m ³ (STP) / h.

The flow divider divides the product gas A at the outlet of the shaft reactor into two parts of the same composition.

The first part was 10.93 m ³ (STP) / h and the second part was 13.36 m ³ (STP) / h. The first portion was recycled to the reaction zone A as a component of the reaction gas A charge gas mixture. For the dehydrogenation circulating gas flow of the first part, the jet pump is operated using the heated compressed residual gas circulating gas as the driving jet, and the delivery of the driving jet compressed through the driving nozzle through the mixing zone and the diffuser Direction indicates the direction of flow into the reaction zone A, the suction nozzle indicates the direction of the first flow portion of the product gas A, and the "suction nozzle-mixing zone-diffuser" connection comprises a first part of the product gas A to be recycled Thereby forming a single connecting line between the approaching portions of the reaction zone A.

The second part of the product gas A is transferred out of the reaction zone A and the indirect residual heat of the indirect gas stream with indirect residual heat of the compressed residual gas circulation gas (10.52 m ³ (STP) / h; T = 135 ° C; P = 3.00 bar) Gt; 330 C &lt; / RTI &gt; by heat exchange (the working pressure is essentially maintained). This heated the residual gas circulation gas to 455 캜 and reduced the working pressure to 2.82 bar.

The product gas A was then cooled to 220 ° C in a second indirect heat exchange with "propane-scrubbed" product gas A (9.06 m ³ (STP) / h; T = 30 ° C., P = 10.50 bar) At the same time, the work pressure was reduced to 2.21 bar.

This "propane-scrubbed" product gas A (hereinafter also referred to as "off-gas") was heated to 300 ° C. and the working pressure was changed to 10.10 bar. Thereafter, the off-gas was compressed from 10.10 bar to 3.50 bar in the first expansion stage of the multistage expansion turbine, and at the same time cooled from 300 ° C to 187 ° C. Then, the off-gas out of the first expansion step was transferred to product gas A at 220 ° C in indirect heat exchange. The latter was essentially cooled to 219 ° C while maintaining the working pressure. This heated the off-gas to 190 DEG C, at which time the working pressure was substantially maintained. The off-gas was then compressed to 1.23 bar in the second expansion stage of the multistage expansion turbine and cooled to 88 占 폚. Subsequently, off-gas was supplied to the residue painting.

The product gas A at 219 ° C and 2.21 bar was then first cooled to 60 ° C in an air-cooled indirect heat exchanger and then cooled to 39 ° C in a chilled water-cooled indirect heat exchanger. The condensate removed during the cooling process was removed by a liquid separator (0.68 kg / h).

The remaining 12.52 m ³ (STP) / h of product gas A was compressed to 4.81 bar in the first compressor stage of the multistage radial compressor. This heated the product gas A at 39 占 폚 to 108 占 폚. Thereafter, the product gas A was cooled to 54 DEG C by air in an indirect heat exchanger, and the condensed water (5.7 g / h) was removed by a liquid separator. In the second compression stage, the remaining 12.51 m ³ (STP) / h of product gas A was compressed to 10.50 bar. This is associated with heating the product gas A to 126 占 폚. The product gas A was first cooled to 54 DEG C by air cooling in an indirect heat exchanger and then cooled to 29 DEG C by indirect heat exchange with cooling water. The condensate removed during the cooling process was removed by a liquid separator (0.28 kg / h). The remaining product gas A of 12.71 m ³ (STP) / h (the working pressure is still essentially 10.50 bar) was transferred to the lower part of the column with random packing. Its content was as follows:

16.0% by volume of propane,

8.60% by volume of propylene,

0.26% by volume of ethane,

0.77% by volume of methane,

H ₂ 4.17% by volume,

O ₂ 0% by volume,

N ₂ 67.48% by volume,

0.39% by volume of H ₂ O,

0 vol% of CO, and

CO ₂ 1.31% by volume.

At the top of the propane scrubbing column, 115.96 kg / h of industrial tetradecane from Haltermann, Germany (as described in II) was introduced at an introduction temperature of 30 ° C (pressure 10.50 bar).

At the top of the scrubbing column, "propane-scrubbed" product gas A (off-gas) was sparged at a pressure of 10.50 bar and a temperature of 30 ° C and a volume of 9.06 m ³ (STP)

0.02% by volume of propane,

0.01% by volume of propylene,

0.22% by volume of ethane,

1.04% by volume of methane,

5.59% by volume of H ₂ ,

O ₂ 0% by volume,

N ₂ 90.60% by volume,

0.39% by volume of H ₂ O,

0 vol% of CO, and

CO ₂ 1.12% by volume.

After indirect heat exchange with the product gas A delivered in reaction zone A as described above and multistage compression to 1.23 bar in a multistage expansion turbine, off-gas was supplied to the residue painting. It may be appropriate to remove the hydrogen molecules present on the industrial scale before compressing offgas. This can be done, for example, by applying a membrane separation as described above to off-gas. As a result, the removed hydrogen molecules can be completely or partially recycled to the reaction zone A (suitably as a component of the reaction gas A fill gas mixture). Alternatively, for example, it can be combusted in a fuel cell.

121.93 kg / h of absorber at a pressure of 10.5 bar and a temperature of 40 ° C were discharged from the bottom of the externally cooled or unheated "propane scrubbing column". It contained 3.14 wt% propane, 1.61 wt% propylene, 0.01 wt% ethane, and 0.09 wt% CO ₂ . Before the adsorbate was delivered to the top of the downstream desorption column, the pressure was compressed to 1.98 bar (for example with an inverted pump or valve; in the case of an inverted pump the mechanical energy released was suitably also the absorbance of the propane- (Used to recompress the liquid effluent of the desorption column). The resulting two-phase mixture was transferred to the top of the desorption column.

A stream of 7.88 m &lt; ³ &gt; (STP) / h compressed to a pressure of 3.00 bar (heated to 175 DEG C) was fed back to the desorption column (likewise with random packing) .

The liquid effluent (115.94 kg / h) of the desorption column was again pumped down to 10.50 bar by the pump and indirectly cooled with cooling water and then recycled to the propane adsorption at the top of the propane scrubbing column (adsorption column) at a temperature of 30 ° C.

At the top of the desorption column, a gas mixture of 10.97 m ³ (STP) / h was sparged at a temperature of 33 ° C and a pressure of 1.98 bar in the following amounts:

17.77% by volume of propane,

9.53% by volume of propylene,

0.11% by volume of ethane,

O ₂ 14.54% by volume,

54.91% by volume of N ₂ ,

1.59% by volume of H ₂ O,

0 vol% of CO, and

0.55% by volume of CO ₂ .

Indirect heat exchange with water vapor increased the temperature of the gas mixture to 140 ° C (which reduced the working pressure to 1.88 bar).

This gas mixture was used as the reaction gas B starting mixture (charging gas mixture) to charge the reaction zone B having the same structure as the above II.

The propylene loading of the catalyst charge in the first oxidation step was selected to be 130 l (STP) / l · h. The salt melt (53 wt.% KNO ₃ , 40 wt.% NaNO ₂ , 7 wt.% NaNO ₃ ) had the following inlet temperature:

T _A = 323 ° C T _B = 327 ° C

T _C = 262 ° C T _D = 265 ° C

Reaction gas B having a temperature of 260 DEG C, a pressure of 1.67 bar and the following amount was spilled out from the intercooler in an amount of 10.99 m &lt; ³ &gt; (STP) / h:

7.98% by volume of acrolein,

0.53% by volume of acrylic acid,

17.74% by volume of propane,

0.56% by volume of propylene,

0.1% by volume of ethane,

O &lt; _{2 &gt;} 3.92 vol%

N ₂ 54.82% by volume,

H ₂ O 11.44% by volume,

0.41% by volume of CO, and

CO ₂ 1.49% by volume.

(Throttling) air 2.21 m ³ (STP) / h compressed to 3 bar (heated to 175 ° C) in the radial compressor, so that the inlet temperature of the reaction gas to the second oxidation stage is 251 (Acrolein loading = 110 l (STP) / l · h).

The product gas mixture (product gas B) in the second oxidation step had a temperature of 270 DEG C and a pressure of 1.55 bar and had the following contents:

0.03% by volume of acrolein,

6.94% by volume of acrylic acid,

0.21% by volume of acetic acid,

15.25% by volume of propane,

0.48% by volume of propylene,

0.09% by volume of ethane,

O ₂ 2.95% by volume,

N ₂ 60.31% by volume,

H ₂ O 10.49% by volume,

0.55% by volume of CO, and

CO ₂ 1.69% by volume.

The product gas B (12.79 m ³ (STP) / h) was fractionally condensed in a tray column as described in WO 2004/035514.

High boiling point water (polyacrylic acid (Michael adduct) etc.) of 13.8 g / h was supplied to the residue painting.

2.82 kg / h of condensed crude acrylic acid having a temperature of 15 캜 and 96.99% by weight of acrylic acid was discharged from the second collecting tray to the tray column on the supply of the product gas B. After addition of a small amount of water, it was subjected to suspension-crystallization as described in WO 2004/035514, the suspension crystals were separated from the mother liquor in a hydraulic scrubbing column and the mother liquor was recycled to the condensation column as described in WO 2004/035514. The purity of the washed suspension crystals is greater than 99.87% by weight of acrylic acid and is suitable for the preparation of water- "superabsorbent" polymers for immediate hygiene applications.

The amount of acidic water condensate discharged into the condensation column from the third collection tray on the feed of product gas B but not recycled to the condensation column was 1.06 kg / h and the temperature was 33 ° C. Likewise, the residue was supplied to the painting.

At the top of the condensation column, a residual gas of 10.52 m ³ (STP) / h was flowed out of the condensation column at a temperature of 33 ° C and a pressure of 1.20 bar and the following contents:

18.36% by volume of propane,

0.49% by volume of propylene,

O ₂ 3.58% by volume,

N ₂ 73.32% by volume,

1.73% by volume of H ₂ O,

0 vol% of CO, and

CO ₂ 1.51% by volume.

The entirety of the residual gas is compressed to 3 bar (heated to 135 DEG C) and, after the indirect heat exchange with the product gas A transferred in reaction zone A as described, (the drive for the jet pump used for the dehydrogenation circulation gas flow Lt; RTI ID = 0.0 &gt; A) &lt; / RTI &gt;

U.S. Provisional Patent Application No. 60/802786 (filed May 24, 2006) is incorporated herein by reference. Many modifications and variations are possible in light of the above teachings. It is therefore contemplated that the invention may be practiced otherwise than as specifically described herein within the scope of the appended claims.

Claims (37)

A) -propane is fed to the first reaction zone A to form a reaction gas A, Transferring the reaction gas A through one or more catalyst layers forming hydrogen molecules and propylene by partial heterogeneously catalyzed dehydrogenation of propane in reaction zone A, - feeding oxygen molecules to reaction zone A, oxidizing at least a portion of the hydrogen molecules present in reaction gas A in reaction zone A with water vapor, - product gas A containing propylene, propane and water vapor is withdrawn from reaction zone A, B) In reaction zone B, product gas A is used in conjunction with the supply of oxygen molecules to charge at least one oxidation reactor with reaction gas B containing propane, propylene and oxygen molecules, and the propylene present therein is inhomogeneously Catalyzed partial gas-phase oxidation to produce acrolein or acrylic acid or mixtures thereof as a target product, and to produce an unconverted propane-containing product gas B, C) delivering the product gas B out of the reaction zone B and removing the target product present therein in the second separation zone II leaving a propane- D) In the separation zone III, the propane present in the residual gas without being recycled to the reaction zone A is absorbed into the organic solvent by absorption from the residual gas to form a propane-containing absorbent, E) In the separation zone IV, propane is removed from the absorption and recycled to the reaction zone A as a propane-containing feed stream, Wherein at least sufficient hydrogen molecules in the reaction zone A are oxidized with water vapor so that the amount of hydrogen oxidized to water vapor in the reaction zone A is 30 to 70 mol% or more of the amount of hydrogen molecules formed in the reaction zone A, A process for the production of acrolein or acrylic acid or mixtures thereof from propane. The process according to claim 1, wherein at least a part of the thermal energy generated by the oxidation of the hydrogen molecules in the reaction zone A is subjected to indirect heat exchange with the reaction gas A or the product gas A or the reaction gas A and the product gas A A method used to heat a gas feed stream fed to Zone A. 3. The method according to claim 1 or 2, wherein the reaction gas B comprises hydrogen molecules. 3. The method according to claim 1 or 2, wherein the reaction gas B does not contain hydrogen molecules. The method according to any one of claims 1 to 5, wherein the product gas A discharged from the reaction zone A is used as it is to charge at least one oxidation reactor in the reaction zone B. 3. A process according to claim 1 or 2, wherein at least 5 mol% of the water vapor present in the product gas (B) is removed by condensation in the separation zone (I). 3. A process according to claim 1 or 2, wherein at least 25 mol% of the water vapor present in the product gas (B) is removed by condensation in the separation zone (I). The process according to any one of the preceding claims, wherein at least 50 mol% of the water vapor present in the product gas (B) is removed by condensation in the separation zone (I). Process according to claim 1 or 2, wherein 5 to 98 mol% of the water vapor present in the product gas B is removed in the separation zone I by condensation. The process according to any one of claims 1 to 3, wherein the working pressure in the reaction zone A is between 1 and 5 bar. Method according to claim 1 or 2, wherein the following relation is applied to the working pressure P in different zones of the process, measured at the entrance to the particular zone in each case: P _{separation zone III} > P _{separation zone IV} > P _{reaction zone A} > P _{separation zone I} > P _{reaction zone B} > P _{separation zone II} . The method according to claim 1 or 2, wherein the hydrogen molecule is burned in the reaction zone A into oxygen molecules before 90 mol% of the total amount of hydrogen molecules formed in the reaction zone A is formed. 3. The method according to claim 1 or 2, wherein hydrogen molecules are burned in the reaction zone A with oxygen molecules before 75 mol% of the total amount of hydrogen molecules formed in the reaction zone A is formed. 3. The method according to claim 1 or 2, wherein hydrogen molecules are fired into oxygen molecules in reaction zone A before 65 mol% of the total amount of hydrogen molecules formed in reaction zone A is formed. Method according to any one of the preceding claims, wherein the water vapor present in the product gas (A) in the separation zone (I) is at least partially removed and this removal is effected by indirect and direct cooling of the product gas (A). 3. A process according to claim 1 or 2, wherein a portion of the residual gas having a composition of the residual gas is recycled to the reaction zone A as a propane-containing feed stream, and wherein the residual gas comprises oxygen molecules. 3. The method according to claim 1 or 2, wherein the oxygen molecules fed to the reaction zone A are also fed as a component of air. 3. A process according to claim 1 or 2, wherein the oxygen molecules fed to the reaction zone A are also fed as a component of a gas comprising not more than 50% by volume of components other than oxygen molecules. Process according to claim 1 or 2, wherein the removal of propane from the absorption in the separation zone IV is carried out by stripping with a vapor-containing gas and the propane-containing stripping gas is recycled to the reaction zone A as a propane- . The process according to claim 1 or 2, wherein the reaction zone A is adiabatically configured. The process according to claim 1 or 2, wherein from 10 to 40 mol% of the total amount of propane fed to reaction zone A based on a single pass through reaction zone A is converted to dehydrogenation in reaction zone A. 3. The process of claim 1 or 2, wherein the reaction zone A comprises at least one tray reactor. The process according to claim 1 or 2, wherein the reaction gas A formed in the reaction zone A has the following content: 50 to 80% by volume of propane, 0.1 to 20% by volume of propylene, H ₂ 0 to 10% by volume, 0-20 vol% N ₂ , and H ₂ O 5 to 15% by volume. The process according to claim 1 or 2, wherein the reaction gas A formed in the reaction zone A has the following content: 55 to 80% by volume of propane, 0.1 to 20% by volume of propylene, 2 to 10% by volume of H ₂ , 1 to 5% by volume of O ₂ , 0-20 vol% N ₂ , and H ₂ O 5 to 15% by volume. The process according to any one of claims 1 to 3, wherein the product gas A discharged from the reaction zone A has the following content: 30 to 50% by volume of propane, 15 to 30% by volume of propylene, H ₂ 0 to 10% by volume, 10 to 25% by volume of H ₂ O, O ₂ 0 to 1 volume%, and N ₂ 0 to 35% by volume. 3. A process according to claim 1 or 2, wherein product gas A is removed in the separation zone IV and cooled by indirect heat exchange with the propane-containing feed stream recycled to the reaction zone A. The process according to any one of claims 1 to 3, wherein the heterogeneously catalyzed partial gas phase oxidation is a two-step partial oxidation of propylene to acrylic acid. 28. The method of claim 27, wherein the reaction gas fed to the second oxidation step has the following content: 3 to 25% by volume of acrolein, 5 to 65% by volume of oxygen molecules, 6 to 70% by volume of propane, 0 to 20% by volume of H ₂ O, 8 to 65% by volume of H ₂ O, and N ₂ 0 to 70% by volume. The method according to claim 1 or 2, wherein the residual gas has the following content: 1 to 20% by volume of H ₂ O, N ₂ 0 to 80% by volume, 10 to 90% by volume of propane, 0 to 20% by volume of H ₂ O, O ₂ 0 to 10% by volume, 1 to 20% by volume of CO ₂ , and CO 0 to 5% by volume. Process according to claim 1 or 2, wherein the propane-containing feedstream removed in the separation zone IV and recycled to the reaction zone A has the following content: 80 to 99.99 mol% propane, 0 to 5 mol% of propylene, and H ₂ O 0 to 20 mol%. A process according to claim 1 or 2, wherein a portion of the residual gas having a composition of the residual gas is recycled to the reaction zone A as a propane-containing feed stream and the residual gas fraction is 10 wt% or less based on the total amount of residual gas . 3. The process according to claim 1 or 2, wherein the oxygen molecules fed to the reaction zone B are also fed as a component of a gas which contains no more than 50% by volume of components other than oxygen molecules. 3. The method according to claim 1 or 2, wherein the oxygen molecules supplied to the reaction zone B are also supplied as components of air. Process according to any one of the preceding claims, characterized in that before step B), in the first separation zone I, the water vapor present in the product gas A is partially or completely condensed by indirect or direct cooling or indirect and direct cooling of the product gas A Thereby leaving product gas A ^* . 3. A process according to claim 1 or 2, wherein, after step C), a portion of the residual gas having a composition of the residual gas is recycled to the reaction zone A as a propane-containing feed stream. The method according to any one of claims 1 to 3, wherein, before step D), any water vapor present in the residual gas that is not recycled to the reaction zone A is partially or completely removed by condensation, Or any of said water vapor and hydrogen molecules are both partially or completely removed by condensation and separation membranes, respectively. delete